1
TWO-STAGE AROMATICS HYDROGENATION
OF BITUMEN-DERIVED LIGHT GAS OIL
A Thesis Submitted to the College of Graduate Studies and Research
In Partial Fulfillment of the Requirements for the
Degree of Master of Science
In the Department of Chemical Engineering
University of Saskatchewan
Saskatoon
By
Abena Owusu-Boakye
© Copyright Abena Owusu-Boakye, August, 2005. All rights reserved
i
COPYRIGHT
The author has agreed to make this thesis freely available to the libraries of University of
Saskatchewan for inspection. Copying of this thesis, either in part or in whole could be
done only with the permission of the professor(s) who supervised this work or in their
absence; permission can be sort from the Head of the Chemical Engineering Department
or the Dean of the College of Graduate Studies. It is also understood that duplication or
any use of this thesis in part and in whole, for financial gain without prior written
approval by the University of Saskatchewan is prohibited. In addition, the author should
be given the due recognition whenever any material in this thesis work is used.
Request for permission to copy to make any other use of the material in this thesis
should be addressed to:
The Head
Department of Chemical Engineering
University of Saskatchewan
57 Campus Drive
Saskatoon, Saskatchewan
S7N 5A9, Canada
ii
ABSTRACT
In this research, two-stage hydrotreating of bitumen-derived light gas oil (LGO)
from Athabasca oil sands was studied. The objective was to catalytically upgrade the
LGO by reducing the aromatics content and enhancing the cetane content via inter-stage
removal of hydrogen sulfide. The impact of hydrogen sulfide inhibition on aromatics
hydrogenation (HDA), hydrodenitrogenation (HDN) and hydrodesulfirization (HDS)
activities was investigated. Experiments for this study were carried out in a trickle-bed
reactor loaded with commercial NiMo/Al2O3 and lab-prepared NiW/Al2O3 in the stage I
and stage II reactors, respectively. Temperature was varied from 350 to 390 oC at the
optimum LHSV and pressure conditions of 0.6 h-1 and 11.0 MPa, respectively. The
results from two-stage process showed significant improvement in HDA, cetane rating
and HDS activities compared to the single-stage process after the inter-stage removal of
hydrogen sulfide. Hence, the presence of hydrogen sulfide in the reaction retarded both
the HDA and HDS processes in the single-stage operation. Negligible hydrogen sulfide
inhibition was however, observed in the HDN process.
Prior to the two-stage hydrotreating study, single-stage hydrotreating reactions
were carried out over commercial NiMo/Al2O3 catalyst to determine the optimum
operating conditions for maximizing hydrogenation of aromatics. A statistical approach
via the Analysis of Variance (ANOVA) technique was used to develop regression
models for predicting the conversion of aromatics, sulfur and nitrogen in the LGO feed.
Experiments were performed at the following operating conditions: temperature (340-
390 oC); pressure (6.9-12.4 MPa) and liquid hourly space velocity, LHSV (0.5-2.0 h-1).
Hydrogen-to-oil ratio was maintained constant at 550 ml/ml. The results showed that the
iii
two-level interaction between temperature and pressure was the only significant
interaction parameter affecting HDA while interaction between temperature and LHSV
was the most important parameter affecting both HDS and HDN activities. A maximum
63 % HDA was obtained at 379 oC, 11.0 MPa and 0.6 h-1. Experiments with NiW/Al2O3
were also performed in a single-stage reactor with LGO blend feedstock by varying
temperature from 340-390 oC at the optimum pressure and space velocity of 11.0 MPa
and 0.6 h-1, respectively. The following order of ease of hydrogenation was observed:
poly- > di- >> monoaromatics. The order of ease of hydrogenation in other LGO
feedstocks (atmospheric light gas oil, ALGO; hydrocrack light gas oil, HLGO; and
vacuum light gas oil, VLGO) was studied and found to follow the order: VLGO >
ALHO > HLGO. Studies on mild hydrocracking (MHC) in the gas oil feedstocks
showed a net increase in gasoline with a corresponding decrease in diesel with
increasing temperature.
Both the single and two-stage HDA and HDS kinetics were modeled using
Langmuir-Hinshelwood rate equations. These models predicted the experimental data
with reasonable accuracy. The degree of conversion of the gas oil fractions in ALGO,
HLGO and VLGO via mild hydrocracking was best described by a pseudo-first order
kinetic model based on a parallel conversion scheme.
iv
ACKNOWLEDGEMENT
I would like to take this opportunity to express my profound gratitude to my supervisors,
Dr. Ajay Kumar Dalai and Dr. John Adjaye for their immense contributions, support and
guidance throughout my master’s program. Special thank you also goes to Dr. Deena
Ferdous and Mr. Christian Botchwey for their assistances: they were ever ready to give
me a helping hand whenever I was faced with a difficult problem. My appreciation also
goes to the members of my committee: Dr. Ding-Yu Peng and Dr. Hui Wang for their
directions, contributions and precious time. Technical assistances from Mr. T. B.
Wellentiny, Mr. Richard Blondin and Mr. Dragan Cekvic are also highly acknowledged.
Financial assistances from NSERC, the University of Saskatchewan Graduate Education
Equity Scholarship and the CRC award to Dr. Dalai are gratefully acknowledged.
Above all, I would like to thank the Almighty God for His divine wisdom, strength and
protection throughout my program.
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DEDICATION
This work is dedicated to my parents,
Mr. and Mrs. Owusu-Boakye,
My brothers
Cyr il, Joel and Kweku Owusu-Boakye
And my fiance
Kwame Koom-Dadzie
vi
TABLE OF CONTENTS
COPYRIGHT i
ABSTRACT ii
ACKNOWLEDGEMENT iv
DEDICATION v
TABLE OF CONTENTS vi
LIST OF TABLES x
LIST OF FIGURES xii
NOMENCLATURE xv
ABBREVIATIONS xviii
1.0 INTRODUCTION 1
1.1 Research background 4
1.2 Knowledge gaps 6
1.3 Hypotheses 6
1.4 Research objectives 7
2.0 LITERATURE REVIEW 9
2.1 Hydrotreating 9
2.2 Hydrogenation of aromatics (HDA) 11
2.3 Aromatic compounds in petroleum fractions 12
2.4 Reaction and thermodynamic properties of HDA 14
2.5 Reactions of sulfur and nitrogen species 16
2.6 Hydrogen sulfide (H2S) inhibition studies 19
vii
2.7 Cetane rating of diesel 20
2.8 Effects of process variables on aromatics hydrogenation 23
2.9 Challenges of aromatics hydrogenation (HAD) 26
2.10 Single-stage hydrogenation of aromatic compounds 27
2.11 Two-stage hydrogenation of aromatic compounds 28
2.12 Hydrogenation catalysts 30
2.12.1 Nature of sulfide catalytic sites 31
2.12.2 Interaction between hydrogenation and hydrogenolysis 32
catalytic sites
2.13 Kinetics of aromatics hydrogenation 36
2.13.1 Power-law kinetic modeling 36
2.13.2 Langmuir-Hinshelwood (L-H) modeling 39
3.0 EXPERIMENTAL 41
3.1 Scope 41
3.2 Statistical design of experiments 41
3.2.1 Test for significance of regression models 43
3.3 Experimental plan 44
3.3.1 Phase I - Single-stage AYHD with sulfidedNiMo/Al2O3 44
3.3.2 Phase II- Single-stage HDA with NiW/Al2O3 45
3.3.3 Phase III- Two-stage hydrotreating of LGO Blend 46
3.3.4 Phase IV- Kinetic modeling 47
3.4 Experimental procedure 47
3.4.1 Catalyst loading 47
viii
3.4.2 Catalyst sulfiding 49
3.4.3 Catalyst activity stabilization 49
3.4.4 Experimental runs 50
3.4.5 Two-stage hydrotreating 51
3.4.6 Deactivation studies 53
3.5 Feed and product analysis 53
4.0 RESULTS AND DISCUSSION 56
4.1 Single-stage HDA over NiMo/Al2O3 57
4.1.1 Statistical analysis 57
4.1.2 Significant interacting parameters affecting HDA 58
4.1.3 Significant interacting parameters affecting HDS 62
and HDN
4.1.4 Impact of temperature and pressure on cetane index 67
4.2 Single-stage hydrotreating with NiW/Al2O3 68
4.2.1 Hydrogenation of aromatics in LGO blend 68
4.2.2 Hydrodesulfurization (HDS) and Hydrodenitrogenation 70
(HDN)
4.2.3 Aromatics hydrogenation of ALGO, HLGO and VLGO 72
4.2.4 Product yield 74
4.3 Two-stage hydrotreating and H2S inhibition studies 75
4.3.1 Impact of H2S removal and LHSV ratio on HDA 78
4.3.2 Impact of H2S removal and LHSV ratio on cetane index 80
4.3.3 Impact of H2S on HDS and HDN 80
ix
4.4 Kinetic studies 85
4.4.1 Single-stage kinetics with NiMo/Al2O3 85
4.4.1.1 Kinetics of HDA 85
4.4.1.2 Kinetics of HDS 87
4.4.1.3 MHC kinetics in ALGO, HLGO and VLGO 89
4.4.2 Two-stage kinetic studies 94
4.4.2.1 Overall HDA and HDS kinetics studies 94
4.4.2.2 Effects of H2S removal on HDA kinetics 95
4.4.2.3 Effects of H2S removal on HDS kinetics 97
4.4.3 Experimental versus model predictions 97
5.0 CONCLUSIONS 102
6.0 RECOMMENDATIONS 103
REFERENCES 104
APPENDIX 114
Appendix A: Experimental calibration 115
Appendix B: Feed and product analysis 120
Appendix C: Log sheets 128
Appendix D: Experimental calculations and mass balance closure 130
Appendix E: Experimental data 133
x
LIST OF TABLES
2.1 Typical ranges of HDA process variables 11
2.2 Typical structures of some aromatic compounds in petroleum 12
distillates
2.3 Aromatic type distribution in untreated light gas oils (LGO) from 13
different sources
2.4 Hydrocarbons and related ignition quality (cetane number) 22
2.5 Network studies on the impact of operating variable on 24-25
aromatics hydrogenation (HDA)
3.1 Actual and coded levels of the design parameters 42
3.2 Design matrix of experimental program for statistical study 42
3.3 Properties of LGO feedstock 50
4.1 Regression models for HDA, HDS and HDN 58
4.2 Effects of temperature and pressure on cetane index 67
4.3 Cetane index improvement in single and two-stage processes at a 81
pressure of 11.0 MPa and total reaction time of 1.67 h
4.4 Kinetic parameters of MHC in ALGO, HLGO and VLGO 93
4.5 Summary of the apparent kinetic parameters of the overall kinetics 96
studies in the single and two-stage processes (temperature: 340 -390 oC;
pressure: 11.0 MPa; total residence time: 1.67 h)
C.1. Sample of the data recording sheet 128
E.1. Total aromatics, sulfur and nitrogen concentrations after the single-stage 132
hydrotreating with commercial NiMo/Al2O3 catalyst
xi
E.2. Aromatics, sulfur and nitrogen concentrations after single-stage 133
hydrotreating over NiMo and NiW
E.3. Simulated distillation data obtained for the feed characterization 134
study of the different light gas oil feedstock
E.4. Mild hydrocracking data of LGO types at a pressure of 135
11.0 MPa and LHSV of 0.6 h-1
E.5. Overall aromatics, sulfur and nitrogen concentrations in the two-stage 136
hydrotreating process
xii
LIST OF FIGURES
1.1 Conventional vs. oil sands production in western Canada: 2
1999-2015 (Canadian Association of Petroleum Producers)
1.2 Key factors affecting aromatics hydrogenation (HDA) 5
2.1 Proposed reaction pathway for hydrogenation of naphthalene 14
at high pressures
2.2 Some organosulfur compounds in petroleum 16
2.3 Reaction pathways in the HDS of dibenzothiophenes 17
2.4 Some nitrogen compounds present in petroleum distillates 17
2.5 Hydrodenitrogenation (HDN) of pyridine 18
2.6 Dual-site mechanism proposed for hydroprocessing of C2H5X 33
over sulfided NiMo HR 346 catalysts at 340 oC and 7 MPa H2
2.7 Transformation of hydrogenation sites into hydrogenolysis sites 34
2.8 Geometric considerations in HDS of dialkylbenzothiophenes 36
3.1 Schematic diagram of catalyst loading in the micro-reactor 48
3.2 Experimental Set- Up(PG-Pressure Gauge; TC-Temperature controller) 52
3.3 Experimental plan for stage I of the two-stage process 54
3.4 Experimental plan for stage II of the two-stage process 55
4.1a Surface response plot for the effect of temperature and pressure 60
on aromatics conversion (LHSV: 1.25 h-1; H2/oil ratio: 550 ml/ml)
4.1b Effect of interaction of temperature and pressure on HDA 61
activity (LHSV: 1.25 h-1; H2/oil ratio: 550 ml/ml)
xiii
4.1c Contour plots for the effects of temperature and pressure on 63
aromatics conversion (LHSV: 1.25 h-1; H2/oil ratio: 550 ml/ml)
4.2a Surface response plots showing the effect of interaction of 65
temperature and LHSV on sulfur conversion (Pressure: 9.5 MPa;
H2/oil ratio: 550 ml/ml)
4.2b Surface response plots: effect of interaction of temperature and LHSV 66
on nitrogen conversion (LHSV: 1.25 h-1; H2/oil ratio: 550 ml/ml)
4.3 Effect of temperature on the rate of hydrogenation of mono, di- 69
and polyaromatics over NiW/Al2O3
4.4 Effect of temperature on the NiW/Al2O3 activity for HDN and HDS 71
(Pressure: 11.0 MPa; LHSV: 0.6 h-1; H2/oil ratio: 550 ml/ml)
4.5 Simulated distillation curves of the VLGO, ALGO and HLGO 73
4.6 Conversion profiles for hydrogenation of total aromatics in ALGO, 76
HLGO and VLGO over NiW/Al2O3 (Pressure: 11.0 MPa; LHSV: 0.6 h-1)
4.7 Effect of temperature and feed type on product yield 77
(Pressure: 11.0 MPa; LHSV: 0.6 h-1)
4.8 Effect of H2S removal on the reaction rate constants of HDA in 79
the two-stage process (Pressure: 11.0MPa, H2/oil ratio: 550ml/ml)
4.9 Effect of H2S removal and LHSV ratio on the overall cetane index 82
in the two-stage process (Pressure: 11.0MPa, H2/oil ratio: 550ml/ml)
4.10 Impact of H2S inhibition on HDS in the two-stage process 84
(Pressure: 11.0MPa; H2/oil ratio: 550ml/ml)
4.11 Arrhenius and Van’ t Hoff plot for single-stage HDA 88
xiv
4.12 Arrhenius and Van’ t Hoff plot for single-stage HDS 90
4.13 Correlated vs. experimental concentrations of total 99
aromatics (CA) at the different LHSV ratios
4.14 Correlated vs. experimental concentrations of product sulfur 100
concentrations (Cs) at the different LHSV ratios
4.15 Correlated vs. experimental concentrations of the heavy gas oil 101
(345+ oC) fractions from the mild hydrocracking data simulated
distillation) in VLGO, ALGO and HLGO
A.1 Calibration curve for mass flow meter 116
A.2 Temperature distribution along the axial length of the reactor 117
A.3 Temperature calibration curve of reactor 118
B.1 Sample 13C-NMR spectra for a hydrotreated sample 120
B.2 TEM micrograph of sulfided NiMo/Al2O3 catalyst 125 B.3 TEM micrograph of sulfided NiW/Al2O3 catalyst 126
xv
NOMENCLATURE
A* Reactant
A*AD Adsorbed reactant species A*
b constant parameter
CA Product concentration of aromatics
CAE Equilibrium concentration of aromatics
CAH Concentration of saturated products
CAO Concentration of aromatics in feed
Car aromatics content [%]
CD Concentration of diesel fraction [wt %]
CG Concentration of gasoline fraction [wt %]
CH Concentration of heavy gas oil fraction
CN Concentration of total nitrogen [wppm]
Cs Product concentration of sulfur species
CSO Concentration of sulfur species in feed
Ei Activation energy [kJ/mol]
f i Fugacity
H Heat of adsorption [kJ/mol]
I Average integrated detector response
I* Inhibitor
Iar integral of total aromatics
Isat integral of total saturates
KA Adsorption equilibrium constant for aromatics [MPa-1]
xvi
kA Rate constant for HDA
kactual Actual rate constant [h-1]
kF Forward rate constant
Kg Gravimetric dilution factor
KH2 Adsorption equilibrium constant for hydrogen [MPa-1]
KH2S Adsorption equilibrium constant for hydrogen sulfide [MPa-1]
kobs Observed rate constant [h-1]
kR Reverse rate constant
Ks Adsorption equilibrium constant of sulfur [MPa-1]
ks Rate constant for HDS
M Mid-boiling point [oC]
M* Mass of test specimen
N Number of experimental runs
n Reaction order
nH2S Number of moles of hydrogen sulfide
NH3 Ammonia
Pa Atmospheric pressure [MPa]
PH2 Partial pressure of hydrogen [MPa]
PH2S Partial pressure of hydrogen sulfide
Po Standard pressure [MPa]
R Universal gas constant
rA Rate of aromatics hydrogenation reaction
rAobs observed reaction rate for hydrogenation of aromatics [% h-1]
rS Rate of hydrodesulfurization reaction
xvii
S Slope
T Temperature [oC]
Ta Actual temperature [oC]
To Standard temperature
Va Actual volume [ml]
Vo Standard volume [ml]
W Lambert W function
x Number of design factors
XA* Fraction of reactant A* adsorbed [%]
XTA Conversion of total aromatics [%]
XTN Conversion of total nitrogen [%]
XTS Conversion of total sulfur [%]
Y y-intercept
YA Mole fractions of aromatic compound
YA Mole fractions of aromatic compound
YAH Mole fractions of saturated aromatic compounds
YAH Mole fractions of saturated aromatic compounds
yexp Experimental response
yp Model prediction response
xviii
ABBREVIATIONS
ALGO Atmospheric light gas oil (160-393 oC)
ANOVA Analysis of Variance
CI Cetane index
CN Cetane number
D Diesel fraction (205-345 oC)
FT Fourier Transform
G Gasoline fraction
H Heavy gas oil fraction
H/C Hydrogen to carbon ratio
H2O Water
H2S Hydrogen sulfide
HC Hydrocarbon
HDA Aromatics hydrogenation
HDN Hydrodenitrogenation
HDO Hydrodeoxygenation
HDS Hydrodesulfurization
HLGO Hydrocrack light gas oil (163-404 oC)
HT Hydrotreating
LGO Light gas oil
LGOB Light gas oil blend (191-420 oC)
LHSV Liquid hourly space velocity [h-1]
MHC Mild hydrocracking
xix
SFC Supercritical fluid chromatigraphy
SSE Sum of squares
TEM Transmission electron microscopy
VLGO Vacuum light gas oil (271-482 oC)
Greek letters
ρ Density measurement (g/cc)
δ− Partial negative
δ+ Partial positive
1
1. INTRODUCTION
The global demand for oil has increased by 150 % over the last 40 years and 20
% in the past two decades to the current 80 million barrels per day and is projected to
grow by 50 % more in the next 20 years (Isaacs, 2004). This demand for oil comes at a
time when there is a gradual decline in supply from relatively cheap conventional crude
and discoveries are not being replaced with new ones (Laherrere, 2003). However, the
world has over twice as much supply of unconventional oil as compared to conventional
oil and it is estimated that there are 8-9 trillion barrels of heavy oil and bitumen in place
worldwide, of which potentially 900 billion barrels of oil are commercially exploitable
with today’s technology (Davis, 2002).
In 2003 the total conventional light and heavy production was 1,120,000 b/d and
by 2015 this is expected to decline to 600,000 b/d. The significant growth in oil sands
production far exceeds the decline in conventional production. Oil sands production
currently make up approximately half of Canada’s total crude oil output and by 2015 it
is expected to account for three quarters of all Western Canadian production (Canadian
Association of Petroleum Producers, 2004-2015 crude oil forecast). Figure 1.1 shows the
conventional petroleum vs. oil sands production in Western Canada alone. Oil sands
production includes both raw bitumen and upgraded synthetic crude oil while the
conventional portion includes light and heavy oil.
2
Figure 1.1: Conventional vs. Oil Sands Production in Western Canada: 1999-2015 (Canadian Association of Petroleum Producers, July 2004)
As production of upgraded oil increases, there is a strong potential for market
limitation for synthetic crude oil (Isaacs, 2004). This is because of the high aromatics
contents of the synthetic crude oil derived from bitumen which consequently reduces
diesel cetane (Wislon and Fisher, 1985). Also, present in these distillates are high
concentrations of sulfur and nitrogen compounds. At the moment, Canadian and United
States refineries are not designed to mix more than 10 to 15 % into their conventional
crude supply to meet end product quality specifications (Isaacs, 2004).
The heightened concern to produce high quality middle distillates has led to the
enforcement of stricter fuel specifications worldwide. New diesel fuel specifications will
mainly require a reduction in sulfur and aromatic content, while the cetane number will
be set to a minimum value of about 53 units (Eliche-Quesada et.al., 2003). Presently, the
3
minimum cetane number should be at least 40 (US EPA, Diesel Fuel Quality, 1999).
Reducing diesel aromatics from 30 to 10 % will reduce NOx emissions by more than 5
%. Cetane number is a measure of the ignition property of the diesel fuel and has an
inverse relationship with the constituent aromatic hydrocarbons. Aromatics, especially
the polycyclics, have very low cetane numbers while paraffin hydrocarbons have
relatively high ones. Thus a key factor for boosting the cetane number is to decrease the
aromatic content in distillates.
Reduction of aromatics has become a key processing parameter in processing
middle distillates and intensive efforts have been made in recent years to develop
catalysts and processes for producing low-aromatic diesel fuel. Several attempts by
researchers including Matarresse et.al., 1983; Wilson and Kriz, 1984 have been made to
optimize the process variables and maximize hydrogenation using existing hydrotreating
catalysts such as NiMo (W) and CoMo supported on γ-alumina. Due to the nature of the
hydrogenation activity of such catalysts, high temperatures are required for
hydrotreating. However, the high temperatures have a negative effect of shifting
equilibrium in favor of the reverse reaction, which is undesirable. To eliminate the
equilibrium effects, high pressures are usually used although it increases the cost of
production and hydrogen consumption (Stanislaus and Cooper, 1996).
Catalytic hydrogenation is an essential process for reducing aromatics since after
removal of sulfur and nitrogen there still remains appreciable amounts of aromatics.
Presence of these aromatics does not only generate particulate emissions but also
decrease the cetane number. Unlike the other hydrotreating reactions;
hydrodesulfurization (HDS) and hydrodenitrogenation (HDN), hydrogenation of
4
aromatics (HDA) is very difficult due to the exothermic and reversible nature of the
reaction. Hence, application of hydrotreating techniques including single and dual stage
processes have been used by many refineries to enhance the catalytic activity for
hydrogenation of aromatics to improve the fuel quality.
1.1 Research background
Product quality is often a significant issue with hydrotreated products from
synthetic crude distillates (Gray, 1994). These fractions can be more aromatic than
conventional crude oil distillate in the same boiling range, which may be a concern in
reducing the total aromatics content of transportation fuel (Pauls and Weight, 1992). To
be able to produce diesel fuel with very low aromatics contents, a thorough
understanding of the effects of process variables, catalyst type and the interaction of
these variables on chemistry and thermodynamic equilibria of different types of aromatic
compounds present in petroleum distillates is necessary. Figure 1.2 shows the functional
relationship of the most important factors affecting hydrogenation of aromatic
compounds.
Generally, it is the feed properties that determine the type of hydroprocessing
technology and catalysts to use for reducing the aromatics content. Real feedstock
contain a complex mixture of hydrocarbons and other elements such as sulfur, nitrogen
and metals which require high reaction temperatures before they can be removed.
Hydrogenation of aromatics on the other hand requires moderate temperatures and high
pressure conditions. Consequently, a tactful combination of the operating conditions
such as temperature, pressure, and space velocity and hydrogen-to-oil ratio is required to
effectively reduce diesel aromatics and remove the other objectionable elements.
5
Figure 1.2: Key factors affecting aromatics hydrogenation (HDA)
Several reactions occur during hydrotreating, some of which are known to affect
reactivity. One such reaction is hydrodesuphurization (HDS) which produces hydrogen
sulfide as one of its by-products. Presence of the hydrogen sulfide is well known to
retard hydrogenation activity of hydrotreating catalysts. (Girgis and Gates, 1991;
Ishihara et.al., 2003; Kabe et.al.,. 1999; Stanislaus and Cooper, 1996). Hence, the choice
of the hydrotreating catalyst plays a vital role in hydrotreating since different catalysts
present various degrees of hydrogenation activities and tolerance to heteroatom
poisoning and/or inhibition in aromatics hydrogenation.
The object of this project is to improve the product quality of a middle distillate
fraction from Athabasca bitumen oil sands. The content of this report includes a careful
review of literature, knowledge gaps and statement of purpose for hydrogenation of
aromatics in petroleum distillates. Also included in the thesis are discussions on
Catalyst Feed
Process
Catalyst Type
- NiMo/alumina
- NiW/Alumina - CoMo/Alumina - Pd/Pt/Zeolites - Support - Active Phases - Promoters
- Reactor/catalyst bed configuration - Catalyst loading/sulfiding - Process conditions
-Pressure drop
- Single-stage
- Dual stage
Composition - Aromatics - Sulfur - Nitrogen
Aromatics Hydrogenation
(HDA)
6
experimental, results, conclusions and recommendations for improving the product
quality of middle distillates from Athabasca oil sands.
1.2 Knowledge gaps
Although a number of studies have been conducted to maximize hydrogenation
of aromatic compounds, most of the data were obtained using model compounds thus
making direct application of the results to industrial processes quite a challenge.
Information gathered from literature so far shows that existing data on the interaction
effects of operating conditions on hydrogenation of aromatics in gas oils from Athabasca
oil sands is scarce. Also, few reports have been published on single and dual stage
kinetic studies of HDA over NiW catalyst supported on alumina in bitumen derived
LGOs from Athabsaca oil sands. Finally, since production of hydrogen-sulfide gas from
sulfur removal is known to inhibit catalytic hydrogenation activity, it is important to
design the hydrotreating process so as to reduce its inhibition effect on hydrogenation.
However, studies so far shows that limited information exist on the kinetics of H2S
inhibition on hydrogenation of aromatics in distillates from synthetic crude using
alumina supported NiMo and NiW catalyst systems in a two-stage process.
1.3 Hypotheses The following hypotheses have been outlined for this research.
1. Although hydrogen partial pressure has been determined to be the key processing
factor affecting aromatics saturation, interactions of temperature, liquid hourly
space velocity (LHSV) and pressure have superior effects on HDA compared to
pressure alone.
7
2. Under the same hydrotreating conditions, HDA characteristics will be different
for each light gas oil feedstock.
3. Clean fuels with relatively low amounts of aromatics and heteroatom can be
successfully produced after a two-stage upgrading as compared to a single-stage
process. Hydrogen-sulfide, produced as a by-product of the HDS process inhibits
hydrogenation of aromatics as well as the HDS reactions during hydrotreating.
Therefore, removing the hydrogen sulfide inter-stage, will improve HDA and
HDS activities.
4. The Langmuir-Hinshelwood rate equation would truly account for the hydrogen
sulfide inhibition in HDA and HDS reactions during hydrotreating.
1.4 Research objectives
The main objective of this research was to study the catalytic upgrading of light
gas oil from Athabasca bitumen by reducing the aromatic contents and thereby
enhancing the diesel cetane using a two-stage hydrotreating process. Within this
objective, various phases of the research were defined with each phase having a set
objective(s):
1. Phase I - Determine the impact of the interaction of temperature, pressure and
liquid hourly space velocity on aromatics hydrogenation and the best
combination of these factors to give maximum HDA, HDS and HDN. This study
will be performed in a single-stage hydrotreater loaded with NiMo/Al2O3
catalyst. A statistical approach will be used to design the experiments and
analyze the hydrotreating data. The intent of this approach is to develop
regression models that describe HDA, HDS and HDN. Another objective of this
8
phase of the experiment is to determine the optimum operating conditions for
maximum HDA.
2. Phase I I - Following the optimum conditions obtained in Phase I, the
hydrogenation activity of NiW/Al2O3 catalyst will be studied on a variety of light
gas oil feedstock from Athabasca bitumen. Experiments for this study will be
performed in a single-stage hydrotreater.
3. Phase I I I - In this phase, the effect of H2S removal on the hydrogenation
propensity of aromatics using NiMo in stage I and NiW in stage II in a two-stage
hydrotreating process unit will be studied.
4. Phase IV- The H2S inhibition kinetics on aromatics hydrogenation and
hydrodesulfurization reactions using Langmuir-Hinshelwood rate equations will
be developed.
9
2.0 L ITERATURE REVIEW
This chapter presents a review on upgrading of middle and heavy gas oil.
Particular emphasis is placed on the concepts of hydrogenation of aromatics and the
challenges involved in hydrogenation of aromatics (HDA) in both model and industrial
feed. The concepts of hydrogenation catalysts and their catalysis as well as kinetic
modeling of hydrogenation of aromatics are also discussed in this chapter.
2.1 Hydrotreating
Hydrotreating is a process to catalytically stabilize petroleum products and or
remove objectionable elements from products or feedstock by reacting them with
hydrogen (Gary and Handwerk, 2001). It is also used to include a variety of catalytic
hydrogenation processes used in fuels refining or for purification of products such as
industrial solvents. In contrast to hydrocracking, hydrotreating produces very little
change in volatility of chemical species (Satterfield, 1981). Hydrotreating also
encompasses processes such as sulfur removal (hydrodesulfurization, HDS), nitrogen
removal (hydrodenitrogenation, HDN) as well as hydrogenation of some or all
unsaturated species including aromatics, present in a feedstock. The typical
hydrotreating catalysts are sulfided CoMo/Al2O3 or NiMo/Al2O3. A minimum
concentration of hydrogen sulfide is usually required to maintain the catalyst in the
sulfided state (Ishihara et.al., 2003; Stanislaus and Cooper, 1996; Girgis and Gates,
1991).
10
During the process of hydrotreating the oil feed is mixed with pure hydrogen
before or after it is preheated to a proper reactor inlet temperature which is usually
below 427 oC to minimize cracking. The combined feed with the hydrogen-rich-gas then
enters the top of the fixed-bed reactor. In the presence of a metal-sulfide catalyst, the
hydrogen reacts with the oil to produce saturated hydrocarbons, hydrogen sulfide,
ammonia, and free metals. The process of saturation of the aromatic rings is aromatics
hydrogenation (HDA). The general reaction mechanism for hydrotreating is shown in
equation 2.1.
Feed Saturated hydrocarbons + H2S + NH3 + H2O (2.1)
Hydrotreating, as an intermediate processing step for catalytic reforming of gas oil
fraction may be carried out for the following reasons (Satterfield, 1981):
1. Prior hydrodesulfurization provides a means for control of air pollution because
some of the sulfur present in a feedstock can be deposited in the form of coke on
the hydrotreating catalyst which would otherwise be emitted to the air.
2. HDN removes nitrogen compounds that otherwise deactivate acidic sites on
hydrotreating catalysts and also contribute to coke formation.
3. Saturation of aromatic rings is required for: cracking of heavier feedstock such as
heavy gas oil. This improves the cetane number of diesel fuels and smoke point
of jet fuel and reduces particulate emission from exhaust gases. Without such
prior saturation, multi-ring aromatic compounds pass through catalytic cracking
reactor and undergo very little or no reaction.
Catalyst, H2
Heat, pressure
11
2.2 Hydrogenation of aromatics (HDA)
Hydrogenation of aromatic compounds (equation 2.2) is reversible and
exothermic with heats of reaction typically in the range of 63-71 kJ/mole (Reid et.al.,
1977 and Jaffe et.al., 1974). Under typical hydrotreating conditions, complete
conversion of aromatics is not possible and as a result, the kinetics is complicated by a
significant reverse reaction at high temperatures.
Aromatics + nH2 Saturated Hydrocarbons (2.2)
High temperatures, low space velocities and high hydrogen partial pressures are
required to achieve appreciable hydrogenation of aromatics. The reactions and
thermodynamic properties of aromatic hydrogenation are further discussed in sections
2.4. Typical ranges of process variables for hydrogenation of aromatics are shown in
Table 2.1.
Table 2.1: Typical ranges of HDA process var iables (Gray, 1994) Process Var iable Range
Temperature 270 - 340 oC
Pressure 0.68 - 20.7 MPa
Hydrogen, per unit of feed for:
recycle 360 m3/m3
consumption 36-142 m3/m3
space velocity (LHSV) 1.5-8.0 h-1
12
2.3 Aromatic compounds in petroleum fractions
Aromatic compounds are a class of organic compounds containing an
unsaturated ring of carbon atoms. Benzene as well as the fused ring systems of
naphthalene, anthracene and their derivatives is included in the aromatic groups.
Analytical techniques such the supercritical fluid chromatography (SFC), high pressure
liquid chromatography (HPLC) and infra-red (IR) techniques (Chasey, 1991; Ijam et.al.,
1990; Wilson et.al., 1985), have detected three main groups of aromatic compounds
present in petroleum fractions. These are the mono, di and polyaromatics. The mono and
diaromatics are predominantly found in middle distillates while the polyaromatics,
constituting three or more fused benzene rings, are found in larger quantities in higher
boiling fraction (> 350 oC). Table 2.2 shows some of the typical aromatic species present
in petroleum fractions.
Table 2.2: Typical structure of some aromatic compounds in petroleum distillates
Type of aromatics Typical structure
Monoaromatic
e.g. alkyl benzene
Diaromatics
e.g. Naphthalene
Triaromatics
e.g.Anthracene
13
There are variations in the amount and type of aromatic species present in petroleum
fractions depending on the origin and processing conditions of the feedstock. Higher
concentrations of aromatics are contained in unconventional crude distillates as
compared to the conventional petroleum crude distillates (Yui, 1989). Table 2.3 shows
the aromatic type distribution in light gas oil fractions from two different sources;
Athabasca (unconventional) and Kuwait petroleum (conventional). Petroleum feedstock
also contain a moderately large concentrations of heteroatom (sulfur and nitrogen),
which are distributed over the whole boiling range and generally increase in
concentration in the higher boiling point fractions.
Table 2.3: Aromatic type distr ibution in untreated light gas oils (LGO) from different sources Source of LGO
Aromatic Type Athabasca* **Kuwait
Mono 20.7 17.7
Di 12.2 11.5
Poly 3.6 4.5
Total Aromatics 36.5 33.7
*unconventional crude distillate * * Conventional crude distillate
14
2.4 Reaction and thermodynamic proper ties of HDA
In the presence of a catalyst and hydrogen gas, aromatic groups are hydrogenated
to give hydroaromatic and naphthenes. The more rings in an aromatic cluster, the more
thermodynamically favorable the hydrogenation reaction. Monoaromatics are the least
reactive and this is due to the unusually high stability of the benzene ring arising from
resonance (Gray, 1994). Poly-aromatics on the other hand, are easily hydrogenated and
can undergo cycles of hydrogenation and dehydrogenation (Figure 2.1). Generally, for
aromatic species containing more than one ring, hydrogenation proceeds via successive
reversible steps and each successive stage requires progressively more vigorous
conditions (higher temperatures and pressures and longer times) for saturation
(Stanislaus and Cooper, 1996).
+
R
+
High-molecular-weight hydrogenolysis/ring-opening/isomerizationproducts
R
H
H
cis-Decalin
trans-Decalin
Low-molecular weightcracking products
Low-molecular weightcracking products
High-molecular weight hydrogenolysis/ring opening/
isomerization producst
TetralinNaphthalene
2H2
3H2
3H2
+ ...
Low-molecular weightcracking products
H
H
Figure 2.1: Proposed reaction pathway for hydrogenation of naphthalene at high pressure (Alber tazzi et.al., 2004).
15
From the reversible and exothermic nature of HDA, Gully et.al., 1963 postulated
that the equilibrium concentration of aromatics (based on equation 2.2) can be
approximated by:
2
1
1 ( )A
nA AH A H
Y
Y Y K P=
+ + × (2.3)
where YA and YAH are the mole fractions of the aromatics and saturated hydrocarbon,
KA is the equilibrium constant and PH2 is the partial pressure of hydrogen. The
equilibrium adsorption coefficient, KA decreases with increasing temperature leading to
a net increase in equilibrium aromatics concentration. Experimental data for calculation
of the equilibrium constants are sparse but the few calculated equilibrium constants
indicate that there is a considerable variation from one family of aromatics to another
(Stanislaus and Cooper, 1996). For example in the hydrogenation of benzene
homologues, the value of the equilibrium constant decreases with an increase in both the
number of side chains and the number of carbon atoms in each side chain (Lepage,1987;
Girgis and Gates,1991). The same is found for naphthalene (Gully et.al., 1963).
The substitution of alkyl groups leads to a very slight decrease in the heat of
hydrogenation. However, for hydrogenation of hydrocarbons on sulfide catalysts
including NiW and NiMo on alumina support, a complete reverse order of reactivity is
observed. In other words, addition of an alkyl group to the aromatic ring favors the
reactivity of these molecules for hydrogenation. This is due to the influence of the π
electron delocalization through resonance on hydrogenation and hydrogenolysis. Thus
hydrogenation is favored by highly electron-donating substituents and it is easier when
the aromatic rings to be hydrogenated are less aromatic (Moreau et.al., 1990).
16
2.5 Reactions of sulfur and nitrogen species
Hydrogenation of the aromatics may involve removal of heteroatom such as
sulfur (HDS) and nitrogen (HDN) species. Sulfur species in petroleum may exist in two
forms: (1) as thiophene and its derivatives in Figure 2.2, which can be resistant to further
processing and (2) as sulfides which are more easily removed.
S S S
DibenzothiopheneBenzothiopheneThiophene
Figure 2.2: Some organosulfur compounds in petroleum (Gray, 1994)
Studies (Gray, 1994) show that the higher order ring compounds are more
reactive than expected and this has been attributed to two main trends: electronic effects
on adsorption of the reactant onto the catalyst and subsequent reaction, and steric
hindrance of substituents. The least reactive sulfur species are the thiophenes. Removal
of sulfur in the presence of a catalyst and hydrogen produces hydrogen sulfide gas (H2S)
as a by-product which can inhibit hydrogenation of aromatic compounds. The two main
pathways by which HDS of thiophenic compounds occur are shown in Figure 2.3. They
are: (1) an initial step of ring hydrogenation followed by sulfur extraction (steps 1 & 2
and steps 1& 4) or (2) direct sulfur extraction-hydrogenolysis (steps 3, 6 and 7).
Depending on the reaction conditions and the type of catalyst used, either pathway can
be favored. Studies by Girgis and Gates, 1991 have shown that hydrotreating with
NiMo/Al2O3 at high hydrogen partial pressures will favor the hydrogenation step.
17
S S
+ H2S
+ H2S
1 2
6
3 4
7
5
Figure 2.3: Reaction pathways in the HDS of dibenzothiophenes (Whitehurst et.al., 1998)
HDN has a direct relationship with hydrogenation of aromatic compounds. This
is because nitrogen is mainly present as heterocyclic aromatic compounds. Two forms of
the heterocyclic nitrogen compounds are found: the non-basic derivatives of pyrole and
indole and the basic derivatives of pyridine (Figure 2.4) (Girgis and Gates, 1991; Ho,
1988).
N
N
H
N N
PyroleIndole Quinoline Acridine
Non-basic nitrogen types Basic nitrogen types
Figure 2.4: Some nitrogen compounds present in petroleum distillates (Gray, 1994)
18
Higher-ring nitrogen species such as acridine and quinoline have also been
identified in gas oils (Schmitter et.al., 1984). The type of nitrogen present is determined
by the species’ structure. Basic nitrogen species contain a six-ringed structure while the
non-basic compounds contain at least one five-ringed member. Unlike HDS reactions,
HDN of heterocyclic compounds follow only the hydrogenation pathway before
nitrogen extraction (hydrogenolysis) (Girgis and Gates, 1991). This is partly because the
C=N bond is very strong compared to the C-H bond (Katzer and Sivasubramanian,
1979; Kabe et.al., 1999). The hydrogenation step thus reduces the large energy of the C-
N bond in the ring thereby enhancing the ease of C-N bond cleavage. This suggests that
HDN is also limited by equilibrium effects. Figure 2.5 illustrates the HDN reaction
pathway in pyridine.
N
N
+ NH3
6H2 2H2
Figure 2.5: Hydrodenitrogenation (HDN) of Pyr idine
High hydrogen partial pressures are usually used in the industry to force
equilibrium towards the products thus, making HDN irreversible. Generally, the
hydrogenation step in HDS is not considered to be critical (Girgis and Gates, 1991;
Mascot, 1982; Variant, 1983) whereas this may pose as a difficult step for HDN (Girgis
and Gates, 1991; Ho, 1988; Perot, 1991).
19
2.6 Hydrogen sulfide (H2S) inhibition studies
Hydrogen sulfide, which is produced as a by-product of sulfur removal, has been
reported to significantly inhibit hydrogenation of aromatic compounds (van Gestal et.al.,
1992; Girgis and Gates, 1991; Massoth et.al., 1982). However, modeling of the
inhibition is complicated since H2S adsorption can modify the catalyst surface leading to
interconversion of catalytic sites and thereby enhancing the hydrogenation activity by
increasing the density of Bronsted acid sites. Some of the studies in an attempt to model
the inhibition are outlined below.
Ishihara et.al., 2003 investigated phenanthrene (PHE) hydrogenation reaction
inhibition over NiMo/Al2O3 induced by the presence of dibenzothiophene (DBT)
molecule in a conventional fixed-bed reactor. Results from their study showed that PHE
hydrogenation in the presence of DBT slightly increased when the DBT concentration
decreased. Furthermore, the progressive reintroduction of DBT in the feed after a
reaction performed in the absence of DBT led to a significant decrease in the PHE
hydrogenation activity as well as DBT conversion. Both the PHE and DBT conversions
exhibited values lower than the initial ones when DBT was reintroduced in the feed.
Their results show that although the presence of sulfur in a feed is essential to preserve
good catalytic performance, some hydrogenation catalytic sites can be permanently
poisoned, thus reducing the hydrogenation activity.
Mild inhibition of biphenyl hydrogenation over CoMo/Al2O3 catalyst due to
hydrogen sulfide has been reported by Satterfield and Gultekin (1981). Kaszetlan et.al.,
1994 studied the influence of H2S partial pressure on the activity of a model MoS2/γ-
Al2O3 catalyst over a wide range of H2S partial pressure from 0-0.3 MPa, under a total
20
pressure of 6 MPa, at 320-410 oC and 0.5-0.7 h-1. Moderate inhibiting effects of H2S on
aromatics hydrogenation were observed on the hydrogenation activity for H2S partial
pressure up to 0.06 MPa. For partial pressures higher than 0.06 MPa, no apparent
inhibiting effect of H2S on the hydrogenation activity was detected. Ancheyta-Juarez
et.al., 1999 also studied the effects of hydrogen sulfide on the hydrotreating of middle
distillates over Co-Mo/Al2O3 catalyst. Using an isothermal fixed-bed reactor, they
discovered that the inhibiting effect of hydrogen sulfide in hydrogenation of aromatics
decreased with increasing temperatures. This is in agreement with other literature
(Gestal et.al., 1992 and Leglise et.al., 1994).
In summary, hydrogen sulfide can modify the catalyst surface (e.g. by increasing
the density of Bronsted acid sites), especially at high concentrations, which occur at
elevated reaction temperatures. Furthermore, hydrogen sulfide is required to maintain
the catalyst in the form of sulfides, rather than oxides. However, excess amounts of the
hydrogen sulfide can lower the rate of hydrogenation and the inhibiting effect of H2S on
the hydrogenation activity varies with the absolute level of H2S in the reactor, and the
reaction conditions.
2.7 Cetane rating of diesel
The ignition properties of diesel fuels are expressed in terms of cetane number
(CN) or cetane index (CI) (analogous to gasoline octane number). Cetane number is the
performance rating of a diesel fuel, corresponding to the percentage of cetane (C16H34)
in a cetane-methylnaphthalene mixture with the same ignition performance. A higher
cetane number indicates greater fuel efficiency. The current minimum cetane index
21
specification is 40 (US EPA, Diesel Fuel Quality, 1999). The cetane index on the other
hand is an estimate of the cetane number.
The cetane number (CN) is measured using a standard diesel test engine
according to ASTM D613 test method and is a function of both the chemical and
physical characteristics of the fuel. Influence of the chemical properties includes the
molecular structures of its constituent hydrocarbons (Wong and Steere, 1982; Gulder
et.al., 1985). For example, a high proportion of normal (unbranched) paraffins (CnH2n+2)
in the fuel, especially those with long molecular chains, generally improves the CN.
However, cycloparaffins and aromatics with their stable structures are more difficult to
break down and ignite, thus reducing the CN.
Since measurement of CN by engine testing requires special equipment as well
as being time consuming and costly, they are estimated using mathematical correlations.
The number derived is called the cetane index and is a function of the physical
properties of the fuel such as the boiling point, aniline point, gravity and density of the
sample (Gary and Handwerk, 2001). The aniline point is the minimum temperature for
complete miscibility of equal volumes of aniline and the test sample under the ASTM
D611 method. A product of high aniline point will be low in aromatics and naphthenes
and, therefore, high in paraffins. One of the correlations (cetane index) for estimating
the cetane number is the ASTM D976 (equation 2.4). Use of this correlation has a
number of limitations; it can only be applied to fuels containing no additives for
boosting CN, they are also applicable to pure hydrocarbons and synthetic fuels although
substantial inaccuracies may occur when used for estimating CI of crude oils, residuals
or products having a volatility of below (260 oC) (500 oF) end points.
22
22 )(log803.97554.074.774416.164174.454 MMCI +×−+×−= ρρ (2.4)
where M, is the mid boiling temperature (oC), and ρ, is density measured at 15 oC.
Apart from the cetane index measuring the performance rating of diesel engines,
carbon monoxide, hydrocarbon and aldehyde emissions depend on the cetane number of
the fuel (Martin et.al., 1997). Table 2.4 shows some hydrocarbons contained in
petroleum fuel and their related ignition qualities.
Table 2.4: Hydrocarbons and related ignition quality (cetane number) Hydrocarbon Empir ical
Formula
Cetane number
CN
Boiling point oC
Paraffins
3-Methylpentane C6H14 30 63.2
n-Heptane C7H16 57 98.4
n-Dodecane C12H26 80 216.2
Cyclohexanes
Cyclohexane C6H12 13 80.8
Methylcyclohexane C7H14 20 100.3
Bicyclohexylhexane C12H22 53 238.5
Benzene and Alkylbenzenes
Benzene C6H6 0 80.1
Toluene C6H8 -5 110.7
n-Amylbenzene C11H16 18 204-5
Naphthalenes
α-Methylnaphthalene C11H10 0 244.8
α-n-Butylnaphthalene C14H16 7 282.5
4−Methyl-4- (β)−naphthylheptane
C18H24 10 136-8
23
2.8 Effects of process var iables on aromatics hydrogenation
The principal operating variables for aromatic hydrogenation are temperature,
hydrogen partial pressure and space velocity. A lot of studies have been performed to
determine the impact of these process variables on aromatics hydrogenation (HDA) and
Table 2.5 provides a summary of some of these studies. For further information on these
studies, the reader is referred to the accompanying references.
For HDS and HDN reactions which occur simultaneously during aromatics
hydrogenation, studies show that increasing temperature and hydrogen pressure will also
increase sulfur and nitrogen removal as well as hydrogen consumption (Gary and
Handwerk, 2001). Whitehurst et.al., 1998 have reported that high hydrogen partial
pressure in HDS processes will lead to a corresponding low hydrogen sulfide partial
pressure, thus reducing the H2S inhibition effects during hydrogenation. Although
increasing temperature improves HDS and HDN activity, excessive temperatures can
cause severe side reactions such as cracking and reforming of saturated components
(Gray, 1994). It can also induce rapid catalyst aging as a result of sintering and coking
(Speight, 2000; Gray, 1994).
24
Table 2.5: Network studies on the impact of operating var iables on aromatics hydrogenation (HDA)
Impact of process
Var iables on HDA
Reactants
Operating Conditions
Results
References
Pressure Temperature
Coal-derived Asphaltenes Coal-derived middle distillates
390,420 450 oC 3 & 10 MPa NiMo/Al2O3
300-400 oC 4-12 MPa
Higher conversions occurred at higher temperatures. Under high H2 pressure for all temperatures carbon aromaticity and number of aromatic ring/structural unit in unreacted asphaltene were lower than those under lower H2 partial pressure. Above 350 oC, rate of hydrogenation and thermodynamic limitation controlled the hydrogenation conversion of aromatics. At partial pressures of H2 around 12 MPa, thermodynamic limitation of HDA was effectively released up to the temperatures of 400 oC
Yoshimoto et.al., 1984 Machida et.al., 1998 Wilson et.al., 1985 Stanislaus and Cooper, 1996
Temperature
Australian coal Coker LGO
300-500 oC Fixed bed reactor, 330-390 oC, 12.4 MPa ,0.5 h-1 NiOMoO3/Al2O3
Hydrogenation of aromatics passes through a maximum at 450 oC. Asphaltenes formed at 500 oC consist of dehydrogenated species produced at lower temperatures. Below 400 oC a small but significant number of carbon atoms are present in alkyl chains. At 350 oC, 0.5 h-1, 12.4 MPa, the coker LGO can be hydrotreated to meet the diesel products specifications. The cetane index and aromatics saturation are both affected by thermodynamic equilibrium at temperatures higher than 370 oC
Charlesworth, 1980 Anabtawi,1993 Wilson et.al., 1985 Ancheyta-Juarez et.al.,1999
25
Table 2.5: Network studies on the impact of operating var iables on aromatics hydrogenation (continued)
Impact of process
Var iables on HDA
Reactants
Operating Conditions
Results
Reference
Temperature Pressure LHSV
Tar (Turow brown coal) Diesel oil Heavy distillates from Woadoan coal
CoMo/Al2O3 NiMo/Al2O3 340-400 oC 4.0-14 MPa 0.5-2.0 h-1 Continuous flow, trickle bed reactor 280-370 oC 2.0-6.0 MPa 1.5-6 h-1
CoMo and NiMo, fixed bed reactor; 350-390 oC 50-150 kg/cm2 G 0.5-2 h-1
Hydrogenation of aromatics is markedly dependent on temperature and pressure In general product density and aromatics decreased as temperature or pressure increased or space velocity decreased. The decrease of aromatics is rapid up to about 340 oC. Beyond this point, the decrease is very slow. The effect of pressure is stronger up to 5 MPa. After this value, the effect is almost negligible. Lower LHSV and higher H2 pressure are much more effective in hydrogenation than the higher reaction temperature up to 390 oC. H2 pressure was most effective to hydrogenate aromatic rings.
Sliwka et.al., 1995 Lappas et.al., 1999. Ancheyta-Juarez et.al.,1999 Sato,1997
26
2.9 Challenges of aromatics hydrogenation (HDA)
Existing middle distillate hydrotreating using conventional catalysts such as
sulfided CoMo, NiMo and NiW supported on alumina has been adapted for reduction of
aromatic compounds in diesel. However, studies (Stanislaus and Cooper, 1996; Yui,
1989; Wilson and Fisher 1985) have shown that complete hydrogenation of aromatics is
not possible due to the thermodynamic equilibrium limitations under typical
hydrotreating conditions. Conventional hydrotreating catalysts designed to reduce sulfur
and nitrogen levels would lower the diesel aromatics only marginally (Cooper and
Donnis, 1996). Thus the composition and properties of distillate products is highly
influenced by the type of catalyst used. Ali et.al. (1997) performed experiments to study
dearomatization, cetane improvement and deep desulfurization of diesel feedstock in a
single-stage reactor. Using three catalysts; CoMo/Al2O3, NiMo/Al2O3, NiW/Al2O3, to
study the hydrogenation activity at reactor temperatures of 325 and 350 oC, hydrogen
pressure of 7.5 MPa and LHSV of 2 h-1 , these workers observed divergent effects of
aromatics content and molecular weight on the cetane index of light cycle oil. Their
results also showed that it was impossible to obtain a diesel product that met stringent
specifications by using one type of catalyst in a single-stage reactor even under severe
operating conditions.
Hydrogenation of aromatics in real feed is inhibited by organic sulfur and
nitrogen compound present in the feed as well as H2S and NH3 produced by HDS and
HDN reactions, respectively. These compounds are strongly adsorbed on the
hydrogenation centers of the hydrotreating catalysts compared to the other centers that
catalyze the essentially hydrogenolysis reaction (Kasztelan and Guillaume, 1994; Girgis
and Gates, 1991). This condition provides a competitive environment for adsorption of
27
nitrogen, sulfur and aromatics compounds in the feed toward the hydrogenation sites.
Preference for adsorption depends largely on the values of adsorption strengths of the
different compounds (Chmielowiec, 1986 and Perot et.al., 1988). The extent of H2S
inhibition on aromatic hydrogenation also depends on the catalyst system under
investigation.
With the increasingly stringent regulations on diesel oil, a lot of attention has
been given to reducing the aromatic contents of distillate fuels. As hydrogenation is
exothermic, hydrogenation of aromatics is favored at low temperatures but conventional
hydrotreating catalysts are only active at high temperatures (Song, 2003). There is
therefore the need to consider hydrogenation at low temperatures (e.g. <300 oC). One of
the best catalysts for low temperature hydrotreating includes noble catalysts. However,
these groups of catalysts have very low resistance to sulfur compounds.
These inhibiting effects together with equilibrium limitations (under normal
operating range of hydrotreating) make aromatics reduction in industrial feeds (e.g.
diesel) more difficult than the other hydrotreating reactions.
2.10 Single–stage hydrogenation of aromatic compounds
Conventional hydrotreating technology (single-stage) is adapted for saturation of
aromatics and it has been recognized that aromatic hydrogenation is more difficult than
sulfur removal under the conditions that are usually used for hydrotreating (Stanislaus
and Cooper, 1996). The existing middle distillate hydrotreaters utilizing the single-stage
process and designed to reduce sulfur and nitrogen levels would lower diesel aromatics
only marginally (Asim and Yoes, 1987; McCulloch et.al., 1987).
28
Similarly, high severities of hydrotreating (high temperature) will introduce
thermodynamic equilibrium limitation into the hydrogenation reaction and reduce the
cetane index of diesel fuel (Lee and de Wind, 1992). To significantly increase the
cetane index, single-stage hydrotreating at high pressures with specially designed
catalysts for hydrogenation, is recommended. However, the specially developed
hydrogenation catalysts such as the supported noble metal catalysts have very low
resistance to sulfur and nitrogen poisoning, which means such catalysts can not be used
for feedstock containing high levels of sulfur and nitrogen.
It is also observed that reduction of total aromatics is much more difficult than
reduction of polyaromatics because saturation of monoaromatics to naphthenes is much
more difficult than saturation of polyaromatics to monoaromatics. Processing feed
blends containing cracked materials, to meet the 10 vol % total aromatics specification,
will require even more severe conditions making the single-stage approach less
economically attractive.
2.11 Two-stage hydrogenation of aromatic compounds
For the existing moderate pressure diesel hydrotreater (single-stage process)
using base-metal catalyst (NiMo or CoMo), reduction in total aromatics content is very
limited, due to the relatively low hydrogenation activity of the base metal catalyst.
Addition of a second stage reactor with a high activity noble metal catalyst can produce
diesel fuel with low aromatics contents. Especially in the case of noble catalysts, a
separate second stage reactor is usually necessary since nitrogen and sulfur- containing
compounds must be removed in the first stage reactor, as they are temporary poisons to
the catalysts.
29
Studies conducted on two-stage hydrotreating include the work by Mahey et.al.,
1992. They used a two-stage hydroprocessing technique to reduce the pronounced
effects of nitrogen-containing species inhibition during hydrocracking of synthetic crude
gas oils. Hydrocracking was performed using NiW catalyst supported on silica-alumina.
Higher gas oil conversions were achieved and the middle distillate product quality was
remarkably improved as the diesel fuel cetane number increased by 13 %. Diesel tests
also indicated that the particulate emissions in exhaust gases were lowered by 20 %.
Chmielowiec (1986) has also demonstrated that product yields can be
remarkably enhanced in a two-stage approach where unconventional crude gas oil is
denitrogenated and then hydrocracked. The two-stage hydroprocessing technique has
also found application in the upgrading of coal-derived liquids where oxygenated
compounds showed on hydrocracking catalysts an inhibiting effect similar to that of
nitrogen compounds (Nishijima et.al.,1987).
Nishijima et.al., 1996 also compared two-stage hydrogenation in coprocessing
oil and light cycle oil (LCO) using both NiMo and NiW on alumina support. The latter
catalyst was used for the second stage upgrading, because sulfur was largely removed in
the first stage hydrogenation over NiMo/Al2O3 catalyst. From their study, a large
improvement in the cetane index (from 36 in the feed to 53 at the end of the second
stage) was observed in the coprocessing oil whereas the cetane improvement in the
(LCO) remained modest (from 30 in the feed to 43 at the end of the second stage).
Generally, the two-stage process has been found to be superior to the single-stage
technology. Especially for feedstock containing a high concentration of polyaromatics,
single-stage hydrotreating will not be efficient for deep hydrogenation. With the two-
stage hydrogenation technique, improved diesel cetane index property can be achieved
30
and fuel quality is also enhanced as sulfur and nitrogen levels are reduced to relatively
low levels in compliance with the stringent legislature on diesel fuels quality. This
technology is being applied in industries world wide (Naber and Stork, 1991; Peries
et.al., 1981; Suchanek, 1990).
2.12 Hydrogenation catalysts
The choice of hydrogenation catalyst is highly dependent on the sulfur and
nitrogen contents in the petroleum feedstock. When hydrotreating is carried out on
feedstock containing appreciable amounts of sulfur and nitrogen compounds, sulfided
NiMo, NiW or CoMo on γ-Al2O3 catalysts are generally used, whereas supported noble
metal catalysts such as platinum or palladium are used for sulfur and nitrogen-free
feedstock. Noble metal catalysts on Y-zeolite supports have increasingly been used for
hydrogenation in light and middle distillates (Suchanek, 1990; Peries et.al., 1991).
Among the Co (Ni)-promoted group VI (Mo or W) metal sulfides on γ-Al2O3, NiW are
widely used to reduce aromatics, sulfur and nitrogen in petroleum feedstock via
hydrotreating. The ranking order for hydrogenation in this group of catalyst is found to
be NiW > NiMo > CoMo (Frank and LePage, 1981).
Most hydrogenation catalysts are used in the reduced and sulfided form prior to
introduction of hydrocarbon feedstock. These catalysts have been described as
consisting of specific stoichiometric combinations of Ni or Co with Mo or W. They exist
as a sulfides containing one Ni or Co atom in combination with two Mo or W atoms,
chemically anchored to the surface of the solid support (generally alumina or silica
alumina) (Whitehurst et.al., 1998). Sulfiding is done by introducing hydrogen sulfide or
31
a low-boiling sulfur- containing compound (liquid feed) onto the catalyst in the presence
of hydrogen. Commonly used sulfur compounds are carbon disulfide, dimethylsulfide,
hydrogen sulfide gas and butanethiol. Sulfiding temperatures are within the range of
180-350 oC at pressures greater than 1.0 MPa (Speight, 2000). For real feed operations,
the commonly used temperatures are 193 and 343 oC at 9.0 MPa.
2.12.1 Nature of sulfide catalytic sites
The location and promotional effects of Co and Ni catalysts have been explained
by several different structural models, such as the monolayer model (Schuit and Gates,
1973; Massoth, 1975), intercalation model (Voorhoeve, 1971; Farragher and Cossee,
1973), contact synergy model (Delmon1979), Co-Mo-S phase model (Topsoe et.al.,
1981, 1986 and1984) and catalytically active Co site model (Duchet et.al., 1983). The
sulfided forms of Co(Ni)-Mo(W) catalysts may be represented as Co9S8 ,MoS2 andWS2.
In the contact synergy model proposed by Delmon (1979 and 1990), it is
assumed that MoS2 and Co9S8 exist as separate crystallites in contact with each other.
The role of the promoter (Co9S8) is to activate and provide hydrogen atoms to MoS2.
The excess hydrogen atoms would then create reduced centers on the MoS2 surface,
which would in effect be the active sites on the catalysts surface.
In the case of the Co-Mo-S (Ni-Mo-S) phase model proposed by Topsoe and co-
workers they explained that the promoter atoms (Co or Ni) are located at the edges of
MoS2-like structures in the plane of Mo cations. Candia et.al., 1984 and Topsoe et.al.,
1986 reported that the relative amount of Co atoms present as Co-Mo-S phase has a
linear correlation with HDS activity. The catalytically active sites for hydrotreating are
viewed as sulfur or anionic vacancies. Some kinetic studies using model compounds
32
have reported the existence of two distinct types of catalytic sites, one responsible for
hydrogenation and the other responsible for hydrogenolysis of heteroatom (Matarresse
et.al., 1983; Massoth and Maralidhar, 1982; Zdrazil, 1988). These two sites have been
used to explain inhibition reactions of hydrogen sulfide during aromatic hydrogenation.
2.12.2 Interaction between hydrogenation and hydrogenolysis catalytic sites
Several workers (Yang and Satterfield, 1983, Gultekin and Satterfield, 1984 and
Girgis and Gates, 1991) have observed inhibition of hydrogenation by H2S during
hydrotreating. Different inhibition effects of H2S on hydrogenation and
hydrodeoxygenation (HDO) of phenols have also been observed by Gevert et.al., 1987.
However, in the presence of aromatic compounds, no inhibiting effect of H2S was
observed on HDS of thiophene (Moreau et.al., 1990) and HDN of 2, 4-dimethyl pyridine
(Ho, et.al., 1984). To explain this result using the two catalytic centers, Geneste and co-
workers (1980 and 1990) conducted a thorough study into hydrogenation of aromatic
compounds and hydrogenolysis of S, N and O-containing model compounds. They
observed that hydrogenation was mainly affected by the aromatic properties of the
molecules and not hydrogenolysis of S, N and O-containing model compounds.
However, hydrogenolysis was found to be dependent on the nature of the heteroatom.
Hence, hydrogenation and hydrogenolysis reactions could proceed by different
adsorption mechanisms; hydrogenation through horizontal π-adsorption and
hydrogenolysis through vertical adsorption by the heteroatom. Following this
observation, the workers proposed a dual-site mechanism (Figure 2.6) involving either
Mo or W atom at different oxidation levels. The higher oxidation state was assigned to
33
hydrogenation and the lower oxidation state was responsible for hydrogenolysis. These
authors then concluded that Mo or W with three sulfur vacancies (higher oxidation) at
the corners are primarily responsible for hydrogenation through π-adsorption and
hydrogenolysis site could be an edge site with two sulfur vacancies.
X..
HydrogenolysisHydrogenation
electron donating siteelectron withdrawing site
Figure 2.6: Dual-site mechanism proposed for hydroprocessing of C2H5X over sulfided NiMo HR 346 catalysts at 340 oC and 7 MPa H2 (Moreau et.al., 1990)
Contrary to the above conclusion, Topsoe, 1989 argued that for a given catalyst,
HDS, HDN and HDA can occur on the same sites. Accordingly, all the hydrotreating
reactions can be visualized as taking place in vacancies in a mixed surface-sulfide-
hydride and nitride phase. Simulating the HDS, HDN and HDA reactions over a series
of sulfided Ni(Co)-Mo hydrotreating catalysts, Topsoe suggested that the major effect
of the Co and Ni promoter atoms is to lower the equilibrium constant for adsorption of
sulfur and nitrogen species.
Another group of researchers have also suggested that hydrogenolysis centers are
derived from the hydrogenation sites on the catalyst surface when H2S is adsorbed
(Figure 2.7). Hence, only one type of sulfur or anion vacancy present on sulfided
catalyst is required. As a result, distribution of type I (promoted sites) and type II (sites
with H2S adsorbed) sites would depend on the sulfidation state of the catalyst and the
34
partial pressure of H2S. Consequently, the effect of H2S on the rate of the hydrotreating
reactions is expected to depend on the H2S partial pressure during reaction (Vivier et.al.,
1991; Topsoe et.al., 1990).
sMo + H2S
S
Mo Hδ+δ− S-H
Figure 2.7: Transformation of hydrogenation sites into hydrogenolysis sites
An illustration of the Type I and Type II sites (crystallites) and 4,6-
dimethyldibenzonthiophene (4,6-DMDBT) molecules in their approximate sizes is
shown in Figure 2.8 (Whitehurst et.al., 1998). The Type I sites are made up of single-
layered crystals with a thickness of about 6 Å. They are found either lying flat or
standing perpendicular to the support surface (Bouwens et.al., 1994). Those found flat
on the support surface are usually limited geometrically, that is, interaction between the
reactant and the catalyst is not in the plane of the MoS2. Thus, access to the active sites
of the catalyst is still limited to approach from only one side (Chianelli, 1984; Daage and
Chianelli, 1994). The reacting molecule can not approach the catalyst perpendicular
since the molecule is much wider that the Type I layer thickness. Type I crystallites
found perpendicular to the support are most likely to occur in crystals bonded to the
alumina surface by Al-O-Mo or Al-S-Mo bonds (Whitehurst et.al., 1998). They allow a
higher probability site access with higher activity.
35
Figure 2.8: Geometr ic considerations in the HDS of dialkyldibenzothiophenes
Unlike the Type I sites, the Type II crystallites (shown in Figure 2.8) occur as
stacks of small crystallites with a height to diameter ratio of approximately 1.5-3
(Bouwens et.al., 1994). They are much more accessible to the reactant and it is noted
that about 75 % of all the edge sites can be approached by 4, 6 DMDBT molecule in a
perpendicular alignment with the alumina surface. Hence, for either geometric or
electronic reasons, the Type I sites are expected to have lower activity than the Type II
sites.
In summary, hydrogenation of aromatics is strongly inhibited by both nitrogen
species and H2S. These compounds are strongly adsorbed on the hydrogenation centers
than on the other centers that catalyze the hydrogenolysis reaction. Consequently, there
36
is competition between different nitrogen, sulfur and aromatic compounds present in the
feed toward adsorption on hydrogenation sites. Preference for adsorption in such a
competitive environment is strongly dependent on the adsorption strengths of the
different compounds (Stanislaus and Cooper, 1996).
2.13 Kinetics of aromatics hydrogenation (HDA)
Several types of model aromatic compounds have been used in the study of HDA
kinetics. Literature information on HDA kinetics in industrial feedstock such as
petroleum and synthetic middle distillates are relatively scarce, as a result of the
complexity of the reactions. Other studies (Ali, 1998; Girgis and Gates, 1991; Van
Gestal et.al.,. 1992; Kasztelan et.al., 1994) have also been conducted to investigate the
inhibition effects of sulfur and nitrogen removal on aromatic hydrogenation. Two main
models used to kinetically model HDA are the power law and Langmuir-Hinshelwood
rate equations (Girgis and Gates, 1991; Sapre and Gates, 1979). The latter accounts for
inhibition reactions while the former is used to represent the overall rate law for the
individual hydrogenation of the various aromatic groups.
2.13.1 Power-law kinetic modeling The equilibrium reaction in equation 2.5 can be used as a basis for developing a
kinetic model for hydrogenation of aromatics in middle distillates.
A + nH2 AH (2.5)
where A is the aromatic species, AH is the saturated hydrocarbon and kf , kr are the rate
constants for the forward and reverse reactions respectively. By assuming that the
forward reaction is pseudo-first order (since hydrogenation is carried out in large
kf
kr
37
hydrogen excess at constant partial pressure) and the reverse reaction is first order in
saturates (Wilson and Kriz 1984), the rate expression is given by:
AA f A r AH
dCr k C k C
dt− = = − (2.6)
where rA is the reaction rate, CA and CAH are the concentrations of the aromatic and
hydrogenated hydrocarbons, respectively.
The equilibrium constant can be applied to the reactants and products when the
reaction goes to completion:
n
Hr
f
nH
nHAH
AA
Pk
k
PAH
A
Pf
fK
222 ][
][ =≈= (2.7)
where KA is the equilibrium constant, f i is the fugacity of either the aromatics or
saturated aromatics and PH2 is the partial pressure of hydrogen. Manipulation and
substitution of equation 2.7 into equation 2.6, the final rate expression can be written as:
AHnHA
f
AfA
A CPK
kCk
dt
dCr
2
−=−=− (2.8)
This model can also be used to analyze the kinetics of aromatics conversion in a tubular
reactor over NiW/γ-Al2O3 at 340-440 oC and 5.0-17.0 MPa hydrogen pressure (Gray,
1994).
Wilson et.al., 1984 developed a similar model for kinetics of hydrogenation of
aromatics in middle distillates by assuming reversible pseudo-first order reaction and
reaction in excess hydrogen gas. On integration of equation 2.6, the following
expression was derived:
ktCC
CC
AEAO
AEA −=−−
ln (2.9)
38
Where CAO is the initial aromatics concentration, CAE is the equilibrium aromatic
concentration, k is the hydrogenation rate constant, and t is the space time.
Studies by these authors on the aromatic reduction of middle-distillate fractions of
Alberta synthetic crude shows that the reaction is kinetically controlled i.e. kf / kr >>1
with the rate expression:
ktC
C
AO
A −=ln (2.10)
Other kinetic models based on a simple first-order reversible kinetics have been
developed. Yui and Sanford (1984) proposed the following rate equation for aromatics
hydrogenation in middle distillates:
)1(2 ArmA
nHf
A CkCPkdt
dC−−=− (2.11)
Where kf and kr are the forward and reverse rate constants, respectively, PH2 is the
hydrogen partial pressure, CA is the concentration of aromatics, n is the reaction order
with respect to hydrogen partial pressure, and m is the reaction order with respect to the
aromatics contents.
Substituting 1/LHSV for t, the final integrated rate equation for equation 2.11 is:
−=−−
LHSV
k
CCC
CC
AEAEAO
AEA 1ln (2.12)
Most of the kinetic studies on model aromatic compounds reported in literature
deal mainly with the reaction pathways and reactivities, rather than quantitative kinetic
models. The very few studies on kinetic modeling have proposed an overall first order
reaction with respect to the aromatic reactant, neglecting the effect of hydrogen partial
pressure and the thermodynamic equilibrium. For example, Sapre and Gates,1981
39
reported a pseudo-first order reaction rate for benzene hydrogenation over presulfided
CoMo/Al2O3 catalyst at 325 oC and 7.1 MPa. A similar study by Lipannas et.al., 1991 on
the kinetics of fluorenes on NiW/Al2O3 catalyst, found the reaction to be first order in
the aromatic hydrocarbon as well.
2.13.2 Langmuir - Hinshelwood (L-H) modeling Most of the kinetic studies using LH inhibition rate equations have been based on
sulfur and nitrogen model compounds with the following general reaction steps;
1. Adsorption of the reactant (A*) on the active site of the catalyst with an
adsorption factor KA*.
2. Reaction of A* on the surface of the catalyst with other reactants adsorbed on
other sites or in the bulk solution to form products
3. Desorption of the products from the active sites into the bulk solution.
In the presence of other species (I), such as H2S which is usually taken as an inhibitor in
competition with aromatics for the same adsorption sites, the concentration of the
adsorbed aromatics reactant [A*AD] is markedly reduced. The rate of conversion of the
reactant (XA*) is thus dependent on the fraction of adsorbed sites covered by the reactant
instead of the actual concentration of the reactant. The observed rate is given by:
** ][ AactualobsobsA XkAkr == (2.13)
and ...)* ][* ][1(
* ][
*
** +++
=IKAK
AKX
IA
AA (2.14)
with the degree of inhibition expressed as:
...)* ][* ][1(
1
* +++ IKAK IA
(2.15)
40
Although the Langmuir-hinshelwood model can represent hydrogenation kinetics better
than the power law model, in general, simple power law models have been used by most
researchers to represent hydrogenation kinetics since the use of Langmuir-Hinshelwood
type of rate equation is complicated and there are too many coefficients that are difficult
to determine.
41
3.0 EXPERIMENTAL
3.1 Scope
This chapter contains a discussion on statistical design of experiments, the
experimental plan and procedure including operation of the fixed bed reactor and
apparatus used to obtain data in this work. A discussion of the feed and product analysis
is also included in this chapter.
3.2 Statistical design of exper iments
Statistical design of experiments was used to design the experimental program
for Phase I of the research. A response surface methodology (RSM) using the central
composite inscribed design (CCI) was adopted. RSM consists of a group of statistical
and mathematical techniques for empirical model building and exploitation that relate an
output or a response to a number of predictors or input variable that affect it (Box et. al.,
1987). Some of the attractive features of using the RSM approach are that conclusions
can be drawn from the initial stage of investigation and are also very effective for
determining the optimal reaction conditions for a given process (Yoon et. al., 1999).
The central composite inscribed (CCI) design was applied with three design
factors or inputs; temperature (T), pressure (P), and liquid hourly space velocity
(LHSV). This design consists of an embedded factorial and fractional factorial design
characterized by central (0), axial (-1, +1) and star (* ) points. The star points represent
the extreme values of each process variable. The coded and actual levels of the design
factors used in this statistical study are shown in Table 3.1 and Table 3.2 shows the
design matrix of the experimental program.
42
Table 3.1: Actual and coded levels of the design parameters
Temperature (oC): 340 350 365 380 390
Pressure (MPa) 6.9 8.2 9.6 11.0 12.4
LHSV (h-1) 0.5 0.8 1.25 1.7 2.0
Codes -* -1 0 +1 +*
Table 3.2: Design matr ix of exper imental program for statistical study Var iables Measured Response: Conversion
Exper imental Run #
Temperature (T) oC
Pressure (P) MPa
LHSV h-1
XTA %
XS %
X TN %
1 365 6.89 1.25 2 390 9.65 1.25 3 340 9.65 1.25 4 365 9.65 1.25 5 365 12.41 1.25 6 350 8.27 1.70 7 380 8.27 1.70 8 350 11.03 1.70
9 380 11.03 1.70 10 365 9.65 2.00 11 350 8.27 0.80 12 380 8.27 0.80 13 350 11.03 0.80 14 380 11.03 0.80 15 365 9.65 0.50 16 365 9.65 1.25 17 365 9.65 1.25 18 365 9.65 1.25 19 365 9.65 1.25 20 365 9.65 1.25
XTA – conversion of total aromatics: XS- conversion of total sulfur: XTN – conversion of total nitrogen
43
The total number of runs (N) required by this design is 20, calculated from;
N = 2x + 2x + 6 = 20 trials (3.1)
where 6 is the number of replicates at the center levels and x is the number of design
factors or input variables under investigation. In comparison to the other factorial
designs, the number of trials needed for a full second order factorial design would be:
N = 3k + 6 = 33 trial (3.2)
A decrease in the total number of trials with the CCI design is significant and the benefit
is more pronounced in the case of 6 factors, where the total number of trials would be 80
and 733 for the central composite and the full second order factorials, respectively
(Rigas et.al., 2000).
The Design Expert software version 6.0 was used to design the experiments and
process the data.
3.2.1 Test for significance of regression models
Analysis of Variance (ANOVA) technique was used to test for the adequacy of
the regression models of HDA, HDS and HDN. This is a test based on the variance
ratios to determine the significant differences among the means of several groups of
responses and their normal distribution. The statistical F-test was used to determine the
significance of effects on the regression models (Yoon et.al., 1999). For any regression
equation to be statistically significant, the probability value (p-value) of the F-values
should be less than 0.05 (User Manual, Design Expert 6.0, 2003).
Other statistical properties such as the R2 coefficient and lack of fit test were used
to check for the goodness of fit of the regression models (Rigas et.al., 2000).
44
Information obtained from the R2 coefficient (which varied from 0-1) determined the
percentage variability of the optimization parameters (conversions) explained by the
model while the lack of fit test was used to validate the selected model. Probability
values, (p-value), of the ‘ lack of fit test’ greater than 0.1 were desired.
3.3 Exper imental plan
3.3.1 Phase I : Single-stage HDA with sulfided NiMo/Al2O3 catalyst
Studies (Wilson et.al., 1985; Yui et. al., 1981; Gary et. al., 2001) show that
temperature, H2 partial pressure, space velocity and hydrogen-to-oil ratio are the main
processing parameters affecting hydrogenation of aromatics. Gary and Handwerk, 2001
reported that among the process factors of hydrotreating, hydrogen partial pressure is the
most important factor affecting HDA. However, most of the existing data in literature
(Wilson et.al., 1985; Yui et. al., 1981; Gary et. al., 2001; Gray, 1994) leading to the
above observations were obtained using the classical one-variable approach to design the
experiments. However, this method of experimental design ignores the interaction
effects of the operating variables on hydrogenation of aromatics and may be inadequate
in determining the optimum conditions for maximizing the response.
The purpose of this part of the thesis was to determine the significant interaction
process variables and the optimum operating conditions of aromatics hydrogenation,
sulfur and nitrogen removal using a statistical technique. Experiments were done using a
blend of LGO feedstock and a commercial NiMo/Al2O3 catalyst in a single-stage
hydrotreater. The response surface methodology via the central composite inscribed
design (CCI) was used to design the experiments and the hydrotreating data were
analyzed by the ANOVA technique.
45
3.3.2 Phase I I : Single-stage HDA over sulfided NiW/Al2O3 catalyst
One of the factors affecting hydrogenation of aromatics is the type of
hydrotreating catalyst used. The common hydrotreating catalysts used in industries are
NiMo, CoMo and NiW on alumina supports. Among these hydrotreating catalysts, the
latter is known to be the most effective for hydrogenation of aromatics followed by
NiMo and CoMo. NiMo is efficient for nitrogen removal but can also be used for some
degree of hydrogenation since majority of nitrogen species found in petroleum are found
attached to aromatics structures.
The purpose of this study was to investigate the activity of NiW/ γ-Al2O3 for
hydrogenation of aromatic compounds in a variety of light gas oil feedstock; vacuum,
atmospheric, hydrocrack and a blend (VLGO, ALGO, HLGO and LGO blend,
respectively). The effects of temperature on the liquid product distribution and mild
hydrocracking (MHC) in the LGO feedstock were also studied. All the experiments in
were performed by varying temperature from 340-390 oC at the optimum pressure and
LHSV conditions obtained in Phase I.
A lab-prepared NiW/Al2O3 was used for hydrotreating the feedstock. The
catalyst was prepared by incipient wetness impregnation method (Ferdous et.al., 2004).
By this approach, a solution containing 3.0 wt % of Ni in nickel nitrate [Ni
(NO3)2.6H2O], 15 wt % of tungsten in ammonium metatungstate and 2.5 wt % of
phosphorus in phosphoric acid (H3PO4) in water was impregnated onto the alumina
support at room temperature. The support (γ-Al2O3, Sud Chemicals India, Ltd., New
Delhi) was initially dried at 120 o C overnight. Following impregnation, the catalyst was
dried at 120 oC for 12 h and then calcined at 500 oC for another 4 hours. The catalyst
46
was characterized for its BET surface area and transmission electron microscopy (TEM)
measurements (see Appendix B).
3.3.3 Phase I I I – Two-stage hydrotreating of LGO Blend
From literature (Stanislaus et.al., 1994; Landau et.al., 1996; Mahay et.al. 1991)
the two-stage hydrotreating process has proven to be more efficient for maximum
hydrogenation of aromatics. Since HDA is deemed more complex than HDS and HDN,
reactions in the stage I reactor are usually targeted at heteroatom removal while the
reactions in the stage II reactor are purposely used for hydrogenation of aromatics
(Nishijima et.al., 1996; Stanislaus and Cooper, 1996).Removal of the heteroatom,
specifically sulfur species in the stage I reactor, is to reduce the overall inhibition of
hydrogen sulfide on HDA.
The focus of this part of the research was to investigate the effect of hydrogen
sulfide removal on the degree of hydrogenation of aromatic compounds in LGO feed
from Athabasca bitumen in a two-stage hydrotreater using two catalyst systems;
NiMo/Al2O3 (stage I) and NiW/Al2O3 (stage II). The effect of residence time on
hydrogenation of aromatics in each stage of hydrotreating was also studied in terms of
the distribution of the liquid hourly space velocity between the stage I and stage II
reactions.
Results from the two-stage process were then compared to those obtained from a
single-stage process where hydrotreating was carried out over commercial NiMo/Al2O3
catalyst. The experiments were performed by varying temperature from 350-390 oC at a
constant pressure of 11.0 MPa and LHSV ratios of 1:1.5; 1:1 and 1.5:1.
47
3.3.4 Phase IV-Kinetic modeling
Kinetic modeling of the single and two-stage experiments was done using the
Langmuir-Hinshelwood rate equations. Data for the kinetic studies were obtained from
the experimental results from Phases I-III. The main objective of this phase of the
research was to develop mathematical models describing inhibition of HAD by H2S
during hydrotreating. Further studies were also conducted on mild hydrocracking
(MHC) during upgrading of the different bitumen-derived LGO feedstock. Kinetics
parameters controlling MHC were determined from power law kinetic models.
3.4 Exper imental procedure
3.4.1 Catalyst loading
The reactor (internal diameter =10 mm and length = 285 mm) was sealed at the
bottom with a Swagelok 60 micron stainless steel filter and then packed with glass
beads, silicon carbide and catalyst material from bottom to top. The extrudate catalyst
(1.2-2.0 mm) was first dried at 200 oC for three hours in an oven before being loaded
into the reactor. A complete catalyst loading was made up of three main parts; separate
sections of various sizes of silicon carbide; catalyst bed and glass beads. Figure 3.1
shows the schematic representation of the catalyst loading in the reactor. Above and
below the catalyst bed are layers of glass beads followed by a 25, 10 and 10 mm of 16,
46 and 60 mesh silicon carbide (SiC), respectively. The catalyst bed is maintained at 10
cm high by diluting the catalyst pellets with 90-mesh size SiC. The purpose of the
diluents as well as the SiC layers is to provide complete catalyst wetting, reduce radial
dispersion and reduce the bed porosity; thus minimizing any diffusion effects and
providing plug flow conditions for isothermal reactions (Bej et.al., 2001).
48
Figure 3.1: Schematic diagram of catalyst loading in the reactor
3.4.2 Catalyst sulfiding
Prior to hydrotreating, the metal-oxide catalyst was transformed to the sulfide
state with a solution of butanethiol in straight-run gas oil. The essence of sulfiding was
to decrease the initial high activity of the catalyst and maintain uniform catalyst activity
across the catalyst surface (Yui, 1994; Nagai et.al., 1988).
Catalyst
3 mm diameter glass beads
16 mesh SiC
90 mesh SiC
60 mesh SiC
46 mesh SiC
Catalyst bed
49
Maintaining the operating pressure at 9.0 MPa, helium was allowed to flow
through the system at 50 ml/min as the reactor temperature was steadily increased to 100
oC. At this temperature, 100 ml of a 2.9 vol % butanethiol solution was pumped into the
catalyst bed at a very high rate to wet the catalyst. The flow rate was then reduced to
maintain an LHSV of 1.0 h-1. Hydrogen gas was then introduced at a rate corresponding
to the hydrogen-to-oil ratio and the helium flow turned off (see Appendix A for a
discussion on the calibration of the hydrogen mass flow meter). The reactor temperature
was then increased to 193 oC. At this condition, sulfiding was allowed to occur for 24
hours. The temperature was then increased to 343 oC for another round of 24-hour
sulfiding.
3.4.3 Catalyst activity stabilization
After sulfiding, the catalyst was stabilized at a temperature of 375 oC, LHSV of
1.0 h-1 and pressure of 9.0 MPa for five days by hydrotreating with heavy gas oil. The
purpose of catalyst stabilization was to ensure uniform catalyst activity across the
catalyst surface prior to the experimental runs (Speight, 2000). Sample products were
collected after every 24 hours, stripped and analyzed for sulfur, nitrogen and aromatics
contents.
3.4.4 Exper imental runs
Four different light gas oil fractions from Athabasca bitumen and produced by
Syncrude Canada Ltd were used for the experimental study. The feeds used were
vacuum light gas oil (VLGO), hydrocrack light gas oil (HLGO), atmospheric light gas
oil (ALGO) and blend of all the light gas oils (BLGO). Although the feedstock are from
50
the same source they have varying aromatics, sulfur and nitrogen content due to the
different processing conditions used for each feedstock. Table 3.3 summarizes the
properties of the feed. Commercial NiMo/Al2O3 catalyst and lab-prepared NiW/Al2O3
catalysts were used for hydrotreating.
Table 3.3: Properties of LGO feedstock Feed 13C-NMR
[%] Total Nitrogen
[wppm] Total Sulfur
[wppm] Cetane index
CI VLGO 23.6 634 26780 41.2
ALGO 15.0 290 15020 36.3
LGO Blend 17.1 461 17420 36.1
HLGO 24.0 1773 7149 43.2
Schematic diagram of the experimental set-up is shown in Figure 3.2. During
hydrotreating, the oil feed was mixed with hydrogen rich gas which entered the top of
the fixed bed reactor in a downward flow pattern. In the presence of the metal sulfide
catalyst, the hydrogen reacted with the oil to produce hydrogen sulfide, ammonia,
saturated hydrocarbons and free metals. The reaction temperature was provided by a
twin-furnace system attached to the reactor and monitored by a temperature controller
(See Appendix A for reactor temperature calibration). The reactor effluent was then
stripped off any ammonia in the scrubber after which it was stored in a high pressure
separator to separate the liquid products from gases. Hydrogen sulfide in the exit gas
was absorbed in a sodium hydroxide solution and the excess hydrogen vented to the
atmosphere. From the high pressure separator, sample products were collected and
stripped off any remaining hydrogen sulfide gas and ammonia by bubbling nitrogen gas
through the sample for at least 2 hours at a slow rate.
51
As shown in Figure 3.2, only one reactor was used throughout the experimental
work. For the two-stage process, the experiment was designed such that all the stage I
experiments were completed before the stage II experiments were performed (see
section 3.4.5 for discussion on the two-stage hydrotreating process).
3.4.5 Two-stage hydrotreating
The LGO blend was used as feedstock for the two-stage upgrading process.
Experiments were performed at 350,365,380 and 390 oC at three different space
velocities ratios between stage I and stage II of 1:1.5, 1:1 and 1.5:1. Pressure was
maintained constant at 11.0 MPa. The combined reaction time for both stages was 1.67 h
corresponding to the same reaction time for hydrotreating of the same feed over
commercial NiMo/Al2O3 in a single-stage hydrotreater (Phase I). The combined reaction
time is also the reciprocal of the optimum LHSV for maximizing hydrogenation as
obtained in Phase I.
The same procedure for catalyst loading was used in both reactors (i.e. 5 g of
NiMo/Al2O3 in the stage I reactor and 5 g of NiW/Al2O3 in stage II reactor). Hydrogen
sulfide was removed in the stage I effluents by bubbling pure nitrogen gas through the
collected product before being further hydrotreated in the stage II reactor. The products
from both stages were tested for total sulfur, nitrogen and aromatics contents. Figure 3.3
and Figure 3.4 show the experimental plan for the two-stage hydrotreating process.
52
Figure 3.2: Exper imental set up (PG-Pressure Gauge; TC- Temperature Controller )
High Pressure S ep a ra t o r ( Pro d uc t S t o ra ge T a n k )
N �
H � S S c rub b er
H � He
T C
PG C hec k V a l v e
R ea c t o r
H �O S c rub b er
PG
PG
T o V en t
B a l a n c e
F eed
B a c k Pressure R egul a t o r
Need l e V a l v e
PG PG PG
PG
F urn a c e
53
3.4.6 Deactivation studies
Deactivation studies were performed at the end of each experiment to determine
the extent of catalyst deactivation. This involved running the experiments at the same
conditions as the catalyst stabilization step for a period of three days. Results from the
deactivation tests were compared to the catalyst activity at the stabilization conditions. A
few deactivation tests were also done at some selected experimental conditions to check
the reproducibility of the data. The deactivation tests did not indicate significant change
in the catalyst activity. A significant loss in activity would indicate that the catalyst
deactivated during the experimental run.
3.5 Feed and product analysis
The feed and products were measured for aromatics, sulfur and nitrogen contents.
The total aromaticity was determined using 13C-NMR spectroscopy while Supercritical
fluid chromatography (SFC) was used to determine the concentrations of the individual
aromatics groups, namely mono, di and polyaromatics. Sulfur and nitrogen
concentrations were measured by combustion/fluorescence or chemiluminescence’s
techniques using an Antek 9000 NS analyzer. Boiling point distribution of the feed and
product samples were analyzed by GC simulated distillation using Varian model CP
3800 gas chromatography. Details of the analytical techniques are given in Appendix B.
The cetane indices (CI) of the feed and sample products were calculated as a function of
density and boiling point temperature using the ASTM D976 correlation.
54
1 d 1 d 5 d 19 d 22 d 1 d 2 d 1 d 5 d 26 d
193 oC
343 oC
375 oC 0.67 h
0.83 h
375 oC
193 oC
343 oC
375 oC
1.0 h
350 oC 356 oC 380 oC 390 oC
350 oC 356 oC 380 oC 390 oC
350 oC 356 oC 380 oC 390 oC
sulfiding
Stabilization
Deactivation studies
Stabilization
5 g of NiMo/Al2O3; Pressure: 11.0 MPa
First stage (Stage I )
Time on stream [Days, d]
Figure 3.3: Exper imental plan for stage I of the two-stage hydrotreating process
Fresh NiMo/Al2O3 catalyst reloaded into the reactor
55
Figure 3.4: Exper imental plan for stage I I of the two-stage hydrotreating process
193 C
343 C
375 C
0. 67 h
0.83 h
1.0 h
375 C
1 d 5 d 8 d 1 d 8 d 8 d 2 d
Time on stream [Days, d]
5 g of NiW/Al2O3, Pressure: 11.0 MPa
Second stage (Stage I I ) D
eactivation studies
Stabilization
Sulfiding
350 oC 356 oC 380 oC 390 oC
350 oC 356 oC 380 oC 390 oC
350 oC 356 oC 380 oC 390 oC
56
4.0 RESULTS AND DISCUSSION
This chapter describes the results obtained at the different phases of the research.
Section 4.1 deals with the study of the significant interaction factors affecting HDA,
HDS and HDN in a single-stage reactor with commercial NiMo catalyst. Optimization
of the HDA process for maximum hydrogenation of aromatics is also discussed in this
section. Studies of the hydrogenation and hydrotreating activities of lab-prepared NiW
catalyst in four different light gas oil feedstocks are discussed in Section 4.2. This
section also gives details on the extent of mild hydrocracking (MHC) in gas oil fractions.
Section 4.3 describes the impact of hydrogen sulfide inhibition on HDA, cetane index
and HDS at three liquid hourly space velocity ratios in a two-stage hydrotreating process
where the commercial NiMo is used in the stage I reactor and the lab-prepared NiW in
the stage II reactor. Finally, Section 4.4 describes the kinetic modeling of HDA, HDS as
well as MHC.
To ensure reproducibility of the results, some of the experiments were repeated.
Measurements of the concentration of aromatics, sulfur and nitrogen concentrations as
well as the simulated distillation showed a maximum variation of 7 wt %, 5 wppm, 2
wppm and 4 wt %, respectively.
57
4.1 Single-stage HDA with NiMo/Al2O3
Statistically designed experiments based on Central Composite Inscribed Design
(CCID) were conducted to investigate the significant interaction parameters controlling
conversion of aromatics, sulfur and nitrogen in HDA, HDS and HDN processes,
respectively. The optimum conditions for maximum hydrogenation of aromatics were
also determined based on the regression model for estimating aromatics conversion. The
experiments were performed by hydrotreating a blend of light gas oil from Athabasca oil
sands over commercial NiMo/Al2O3. The operating conditions used were: temperature
(340-390 oC); pressure (6.9-12.4 MPa) and LHSV (0.5-2.0 h-1). Hydrogen-to-oil ratio
was maintained constant at 550 ml/ml.
4.1.1 Statistical analysis
The Analysis of Variance (ANOVA) technique was used to develop response
surface, interaction and contour plots as well as regression models for predicting the
percent conversions of aromatics, sulfur and nitrogen in HDA, HDS and HDN
processes, respectively.
The following linear regression model consisting of the main effects, interactions
and quadratic terms, was used (Box et.al., 1978).
233
1iiij
jiiiji
iio XXXXY ∑∑∑ +++=
<=
αααα (4.1)
where Y is the estimate of the response variable and X i’s are the independent variables
(temperature, pressure and LHSV) for each experimental run. The expressions αo,
αi αij and αιι are the regression parameters. The main effects are represented by the
X i’s, X iX j’s account for the interaction terms, and X i2 terms indicate quadratic effects.
58
Conversion is defined as:
%100][
][][ ×−=feed
productsfeedConversion (4.2)
where [feed] and [products] are the concentrations of the species in the feed and product
samples, respectively.
4.1.2 Significant interacting parameters affecting HDA
The HDA model in Table 4.1 shows that the two-level interaction between
temperature and pressure is the only significant interaction term influencing
hydrogenation of aromatic compounds in the bitumen-derived light gas oil.
Table 4.1: Regression models for HDA, HDS and HDN
* Interaction terms Y i = regression model
Parameter Model coefficients for estimating conversion
Model YHDA YHDS YHDN
Intercept 50.10 97.80 97.60
T 15.20 2.60 4.50
P 21.80 -0.90 0.95
LHSV -6.70 -2.50 -5.90
T2 -26.30 -1.80 -4.60
P2 -41.10 <0.01 -2.10
LHSV2 <0.01 -2.80 -4.40
T × P* 23.60 <0.01 <0.01
T ×LHSV* <0.01 1.50 4.70
LHSV × P* <0.01 <0.01 <0.01
59
This means that at any constant value of LHSV, a simultaneous increase in both
temperature and pressure will significantly increase conversion of the aromatics.
The three-dimensional plot in Figure 4.1a shows that conversion of aromatics
passes through a maximum with increasing temperature and pressure. That is, increasing
temperature accelerates the rate of reaction until the thermodynamic equilibrium
limitation begins to exert a significant reverse effect on the hydrogenation reaction. The
thermodynamic effect is as a result of the exothermic nature of the reversible reaction
which shifts equilibrium to the reactants, thus producing more aromatics in the products.
The effect of pressure on aromatics conversion is further illustrated in the two-
dimensional plot of Figure 4.1b. This is an interaction plot of the effect of temperature
and pressure (at the two extreme levels) on the conversion profile of aromatics at a
constant LHSV of 1.25h-1. Conversion of aromatics is observed to pass through a
maximum with increasing temperature and pressure. At lower pressure levels (6.9 MPa)
less aromatic compounds are hydrogenated but when the reactor pressure is increased to
12.4 MPa, the hydrogenation activity increases significantly with higher conversions.
This is because when the reactor pressure is increased, equilibrium is essentially forced
towards the products, thus making hydrogenation virtually irreversible with a resulting
increase in conversion (Gray, 1994). It can be inferred from Figure 4.1b that although
higher pressures increase the overall conversion of aromatics, the reaction is still
dominated by equilibrium effects at higher reaction temperatures.
The predicted model gives numerical values of the effects of the input variables
on the response. However, it is difficult to see right away the dependence of the response
surface on the design factors and to be able to achieve this, contour plots are usually
used.
60
Figure 4.1a: Sur face response plot for the effect of temperature and pressure on aromatics conversion (LHSV=1.25 h-1, H2/oil ratio =550 ml/ml)
61
F
igur
e 4.
1b:
Eff
ect
of in
tera
ctio
n of
tem
pera
ture
and
pre
ssur
e on
HD
A a
ctiv
ity
(LH
SV=1
.25
h-1, H
2/oi
l rat
io =
550
ml/m
l)
12.4
MP
a 6.
9 M
Pa
62
Contours are 2-dimensional plots which give a geometric representation of the
underlying response function over the experimental region. Figure 4.1c is a contour plot
where each contour line represents the predicted value of response at a constant LHSV
slice of 1.25 h-1. As predicted by the model, it can be observed from Figure 4.1c, that the
conversion of aromatics covers the range of -12.0 to 57.8 % when both temperature and
pressure change from 340 to 390 oC and 6.9 to 12.4 MPa respectively. The negative
conversion values observed at the extreme levels of temperature and pressure in the
contour plot are due to the low temperature hydrogenation effects and equilibrium
limitations at higher temperatures: at these conditions more aromatics are collected in
the hydrotreated products compared to the feed thus, leading to negative conversion
values.
Optimization of the aromatics hydrogenation shows that at the following
operating conditions; temperature of 379 oC, pressure of 11.0 MPa and LHSV of 0.6 h-1,
conversion of aromatics can be maximized to 63 %. At these same conditions, sulfur and
nitrogen conversions are 98.5 and 99.7 %, respectively. This result suggests that in order
to maximize hydrogenation of aromatics in light gas oil feedstock from Athabasca
bitumen, severe hydrotreating conditions are necessary.
4.1.3 Significant interacting parameters affecting HDS and HDN
The final models for HDS and HDN activities are also shown in Table 4.1. It can
be inferred from both models that the interaction between temperature and LHSV was
the most significant term affecting conversion. Thus, a simultaneous increase in
temperature and space velocity at any constant pressure level would increase conversion
of sulfur and nitrogen during hydrotreating.
63
Figure 4.1c: Contour plots for the effects of temperature and pressure on aromatics conversion (LHSV=1.25 h-1, H2/oil ratio =550 ml/ml)
64
Unlike HDN, increasing pressure for HDS with all other parameters remaining constant
will have a negative impact on the sulfur conversion. This may be due to the fact that the
reaction is already taking place in excess hydrogen therefore any further increase in
pressure will have little or no effect on the system. Yui et.al., 1988 studied
hydrogenation of coker naphtha with NiMo catalyst with the following reaction
conditions; temperature (140-280 oC), pressure (3.0-5.0 MPa) and LHSV of (1.0-2.0 h-1).
Although their conditions are lower than the conditions used in this study, they also
observed that during hydrotreating, temperature and LHSV are the only factors
exhibiting remarkable effect on HDS- pressure had negligible influence on the HDS
activity.
The response surfaces for HDS and HDN in Figures 4.2a and 4.2b, respectively,
show high conversions of sulfur and nitrogen during hydrotreating. Sulfur conversion
varied from ~88 to 99 wt % while nitrogen conversion approached 100 wt %. However,
in both cases sulfur and nitrogen conversions passed through maximum with increasing
temperature and pressure. This is due to equilibrium limitations affecting the reactions
which mean that for sulfur and nitrogen species present in the feed, heteroatoms
preferably react by hydrogenation followed by C-S and C-N bond cleavage due to the
high hydrogenation activity of NiMo catalyst (Massoth et.al., 1990; Knudsen et.al.,
1999).
65
Fig
ure
4.2a
: Su
rfac
e re
spon
se p
lots
sho
win
g th
e ef
fect
of
inte
ract
ion
of te
mpe
ratu
re a
nd L
HSV
on
sulf
ur
conv
ersi
on (
HD
S ac
tivi
ty)
(LH
SV=1
.25
h-1;
H2/
oil r
atio
= 5
50 m
l/ml)
66
Fig
ure
4.2b
: Su
rfac
e re
spon
se p
lots
sho
win
g th
e ef
fect
of
inte
ract
ion
of t
empe
ratu
re a
nd L
HSV
on
nitr
ogen
con
vers
ion
(HD
N a
ctiv
ity)
(L
HSV
: 1.
25 h
-1;
H2/
oil r
atio
: 55
0 m
l/ml)
67
4.1.4 Impact of temperature and pressure on cetane index (CI )
The effect of temperature and pressure on the CI was studied by varying
temperature and pressure between 340-390 oC and 6.9-12.4 MPa respectively, at a
constant LHSV of 1.25 h-1. The cetane index of the feed and hydrotreated samples was
calculated from the ASTM D976 correlation:
22 )(log803.97554.074.774416.164174.454 MMCI +−+−= ρρ (4.3)
where M = mid boiling point temperature (oC) and ρ = specific gravity.
Table 4.2 shows the effect of temperature and pressure on the cetane index.
Table 4.2: Effect of temperature and pressure on cetane index Impact of temperature on the cetane index at a constant pressure of 9.6 MPa
Feed 36
Temperature [oC] 340 365 390
Cetane index,CI 42 ± 0.6 45 ± 0.7 34 ± 0.5
Impact of pressure on cetane index at a constant temperature of 365 oC
Pressure [MPa] 6.9 9.6 12.4
Cetane index,CI 41 ± 0.6 45 ± 0.7 47 ±0.7
It is observed that increasing temperature from 340 to 365 oC leads to a marginal
increase in cetane index from 42 to 45 after which further rise in temperature results in a
progressive decrease in CI to about 34 at 390 oC. Hence, the cetane index of the diesel
fraction also passes through a maximum and this is because of the direct relationship of
the cetane with changes in the aromatics contents during hydrotreating. No equilibrium
effects were however observed when the reaction pressure was varied from 6.9 to 12.4
MPa at a constant temperature of 365 oC and space velocity of 1.25 h-1. The cetane index
68
rather increased from 41 to 47, well above the minimum specification of 40 (US EPA,
1999).
4.2 PHASE I I - Single-stage hydrotreating with NiW/Al2O3
4.2.1 Hydrogenation of aromatics (HDA) in LGO blend
Following the Phase I study, another set of experiments was conducted to
determine the activity of NiW/Al2O3 for hydrogenation of the total aromatics in different
LGO feedstock. Prior to this study, hydrogenation of the mono, di and polyaromatics
contained in a blend of light gas oils from Athabasca bitumen was investigated by
varying the reaction temperature from 340-390 oC at the optimum pressure and LHSV
conditions of 11.0 MPa and 0.6 h-1, respectively.
Figure 4.3 shows the effect of temperature on the rate constants of mono-, di-
and polyaromatic hydrogenation. The reaction rate constants were used as a measure of
the speed of disappearance of poly, di or monoaromatics species during hydrogenation.
The rate constants were derived from the pseudo-first order power law relation:
−=io
ii C
CLHSVk ln (4.4)
where k is the rate constant, LHSV is the liquid hourly space velocity, Ci is the
concentration of aromatics in the products, and Cio is the concentration of aromatics in
the feed. The subscript ‘ i’ refers to the mono-, di- or polyaromatics species.
The rate of disappearance of the aromatic groups ranged from 0.63×10-4 s-1 for
monoaromatics to 2.4×10-4 s-1 for polyaromatics. In terms of kinetics, the ease of
hydrogenation followed the general order: polyaromatics > diaromatics >>
monoaromatics with the fastest step (hydrogenation of polyaromatics) being about 6
69
orders of magnitude greater than the slowest step (hydrogenation of monoaromatics).
Although the reactivity of a compound decreases with increasing molecular weight, the
ease of hydrogenation of the higher order aromatic compounds (poly- and diaromatics)
compared to the monoaromatics is because hydrogenation of the higher order aromatic
species are thermodynamically favored even under mild hydrotreating conditions
(Stanislaus and Cooper, 1996). The lower hydrogenation rate of the monoaromatics is
because of the species’ extra stability provided by its resonance structure.
0
3
6
9
12
330 340 350 360 370 380 390 400Temper ture [oC]
Rat
e co
nsta
nt, k
[10
-4 s
-1]
mono di poly
Figure 4.3: Effect of temperature on the rate of hydrogenation of mono, di and polyaromatics over NiW/Al2O3
70
It can be inferred from this result that hydrogenation of the monoaromatics is the
most difficult and the rate-limiting step and this is consistent with other reports by
Stanislaus and Cooper, 1996 and McCulloch, 1975. Hence, severe hydrotreating
conditions are required to produce high quality diesel fuel since a significant increase in
cetane number is observed when monoaromatics are fully hydrogenated (Hill, et.al.,
2002).
4.2.2 Hydrodesulfur ization (HDS) and Hydrodenitrogenation (HDN)
The effect of temperature on the rate of HDN and HDS over the NiW catalyst is
shown in Figure 4.4. The rate constant values for the two processes were similar at lower
temperatures, i.e. 340-350 oC. However, between 365-390 oC, the HDN activity was faster
than that of the HDS but both reactions approached equilibrium with the former being
about 1.5 orders of magnitude higher. Evaluation of the conversion data also showed that
at equilibrium, nitrogen and sulfur conversions were approximately 99 and 96 %,
respectively. The high HDN activity is due to the high hydrogenation propensity of the
NiW/Al2O3 catalyst since most of the nitrogen species found in LGO have aromatic
structures and removal of nitrogen proceeds by hydrogenation before hydrogenolysis.
Another explanation to the higher HDN activity is the presence of less refractory nitrogen
compounds in the feedstock (451 wppm) compared to the more refractory sulfur-
containing compounds (18,451 wppm). Similar results have been reported by Botchwey
et.al., 2003, when they also studied the effect of temperature on HDS and HDN activities
in bitumen-derived heavy gas oil over NiMo/Al2O3 at the following operating conditions:
temperature of 340-420 oC; pressure of 6.5-11.0 MPa; LHSV of 0.5-2.0 h-1 and H2/oil ratio
of 600ml/ml.
71
0
2
4
6
8
10
330 340 350 360 370 380 390 400
Temperature [oC]
Rat
e co
nsta
nts
[10
-4s-1
]
ks kn
Figure 4.4: Effect of temperature on the NiW/Al2O3 activity for HDN and HDS (Pressure of 11.0 MPa, LHSV of 0.6 h-1; H2/oil ratio of 550 ml/ml)
72
4.2.3 HDA of ALGO, HLGO and VLGO feedstock
The NiW/Al2O3 catalyst was also used to hydrotreat three other light gas oil
feedstock types to determine their difference in hydrogenation reactivity and product
yield. The feedstock used were atmospheric light gas oil, ALGO (160-393 oC);
hydrocrack light gas oil, HLGO (163-404 oC) and vacuum light gas oil, VLGO (271-482
oC). The feedstocks have different compositions due to the varying processing
conditions; ALGO is produced under atmospheric pressure conditions in a distillation
unit. Bottoms from the atmospheric distillation unit are then sent to a vacuum distillation
plant where VLGO is produced. Vacuum is needed for the production of the VLGO so
as to lower the boiling temperature of the material, thereby allowing distillation without
excessive decomposition. HLGO is produced under high pressure conditions via
catalytic hydrogenation with a very high hydrogen/carbon (H/C) ratio.
The feedstock were first characterized to determine the distributions of boiling
temperatures as a function of the amount and type of aromatics, sulfur and nitrogen
contents. Simulated distillation curves of the LGO feedstock are shown in Figure 4.5. It
is obvious that VLGO contains the highest boiling, possibly more complex and less
reactive species followed by the ALGO and HLGO. Differences in the distribution of
the boiling temperatures are mainly due to the sulfur contents in the feed (Ancheyta
et.al., 2004 and Botchwey et.al., 2003). Above the 50 wt % fraction distilled, the
distillation curves for both the ALGO and HLGO are the same, indicating the presence
of similar sulfur species in both feeds. The dominant aromatic group in the ALGO
feedstock may be diaromatics while the HLGO may also be dominated by
monoaromatics due to the pretreatment of the feed.
73
Figure 4.5: Simulated distillation curves of the VLGO, ALGO and HLGO
0
150
300
450
600
0 25 50 75 100
Fraction distilled off [wt %]
Boi
ling
Tem
pera
ture
[ o C
]
VLGO
ALGO
HLGO
74
Figure 4.6 shows the plot of percent saturation of total aromatics in the LGO
feedstock as a function of temperature. The order of ease of hydrogenation was VLGO
>ALGO > HLGO. Since hydrogenation of higher order aromatic compounds is more
thermodynamically favored than the lower group aromatics (Stanislaus and Cooper,
1996), the observed trend suggests that the dominant aromatic groups in the VLGO are
polyaromatics whereas the ALGO and HLGO are dominated by di- and monoaromatics,
respectively. Therefore, for every mole of polyaromatics that is saturated, a mole is
added to the diaromatics group and for each mole of diaromatics reacted; a mole of
monoaromatics is added to the existing monoaromatics content (McCulloch, 1975). It
can be concluded that polyaromatics react more easily than the diaromatics, which in
turn undergoes faster hydrogenation than the monoaromatics.
4.2.4 Product yield
The feed and products of the LGO feedstock were grouped into three main cuts
based on their boiling point distributions: gasoline, (40–205 ºC); diesel (205-345 ºC) and
the heavy gas oil (345+ ºC). Although the simulated distillation data shows that the
feedstock are predominantly light gas oil, some tail-end cuts including heavy and
vacuum gas oil fractions are also present. In this study all fractions above the 345+ oC
were grouped as heavy gas oil. Product yield, in terms of gasoline and diesel production,
was defined as:
Product yield in the various LGO feestocks as a function of temperature is illustrated in
Figure 4.7. When the feedstock were subjected to the same hydrotreating conditions,
Yield = Products (desired)
Reactants × 100 % (4.5)
75
there was a net increase in gasoline yield with a corresponding net decrease in diesel for
all feed types except VLGO. The Hydrocrack light gas oil feedstock gave the highest
gasoline yield at 365 oC. Maximum gasoline yield in the ALGO was 27 % at 390 oC. In
the case of the VLGO there was a net steady increase in both the gasoline and the diesel
fractions with increasing temperature. However, the diesel production was higher
compared to the gasoline production.
4.3 Two-stage hydrotreating and H2S inhibition studies
Several classes of reactions occur simultaneously in hydrotreating–HDA, HDS
and HDN -and the presence of some of the reactants and products are known to
markedly affect the reactivity (Girgis and Gates, 1991). In particular, H2S produced from
HDS reactions have been reported to be responsible for the hindrance of HDA as well as
HDS and HDN reactions (Ishihara et.al., 2003; Kabe et.al., 1999; Nagai et.al., 1998).
In this part of the research the effect of hydrogen sulfide on hydrogenation,
cetane index improvement, hydrodesulfurization and hydrodenitrogenation was studied
at different space velocity ratios (ratio of the LHSV between stage I and stage II
reactors) and temperatures using a two-stage hydrotreating unit with NiMo/Al2O3 in
stage I and NiW/Al2O3 in stage II. The three LHSV ratio distributions between stages I
and II are 1.5: 1, 1:1 and 1:1.5. At each set of LHSV ratio, temperature was varied from
350-390 oC. The combined reaction time of hydrotreating was, however, the same for all
sets of experiments. The results from two-stage unit were then compared to those from
the single-stage process where hydrotreating was carried out over commercial
NiMo/Al2O3 catalyst.
76
0
20
40
60
80
100
330 340 350 360 370 380 390 400Temperature [oC]
Con
vers
ion
of 13
C-
NM
R a
rom
atic
s [%
]ALGO HLGO VLGO
Figure 4.6: Conversion profiles for hydrogenation of total aromatics in ALGO, HLGO and VLGO over NiW/Al2O3 (Pressure: 11.0 MPa; LHSV: 0.6 h-1)
77
010203040
340
350
365
380
390
Tem
pera
ture
[o C
]
Yield [%]
VL
GO
gas
olin
eA
LG
O g
asol
ine
HL
GO
gas
olin
eV
LG
O d
iese
l
F
igur
e 4.
7: E
ffec
t of
tem
pera
ture
and
fee
d ty
pe o
n pr
oduc
t yi
eld
(Pre
ssur
e: 1
1.0
MP
a an
d L
HSV
: 0.
6 h-1
)
78
4.3.1 Impact of H2S removal and LHSV ratio on HDA
Figure 4.8 illustrates the relative gain in HDA rate constants at the different
LHSV ratios in the two-stage process compared to the single- stage process. Relative
gain is defined as:
where ki is the reaction rate constant for HDA. The rate constants were derived from the
Langmuir-Hinshelwood rate equations. Discussion of the determination of the kinetic
parameters is presented in Section 4.4.
A general decrease in rate constants, kA was observed with increasing
temperature and this is because of the enhancement in the hydrogenation activity with
increasing temperature and inter-stage removal of hydrogen sulfide. The negative gain in
the reaction rate constants at higher temperatures is an indication that removal of H2S at
higher temperatures is not significant to the HDA activity since any hydrogen sulfide
produced at these temperatures are quickly desorbed from the surface of the catalyst.
The best activity for HDA in the two-stage was observed for the reaction with LHSV
ratio of 1.5:1.0 between stage I and stage II. This is because of the higher hydrogenation
activity and the longer reaction time on the NiW/Al2O3 catalyst in the stage II reactor.
Hence, more NiW/Al2O3 catalysts may be loaded into the stage II reactor to maximize
HDA activity.
Relative gain, ki = k single-stage
k two-stage – k single-stage (4.6)
79
-10
0
10
20
30
40
50
60
70
350 365 380 390
Temperature [oC]
Rel
ativ
e ga
in in
rat
e co
nsta
nt, k
A[
%]1. : 1.5 1. : 1. 1.5 : 1
Figure 4.8: Effect of H2S removal on the reaction rate constants of HDA in the two-stage process (Pressure: 11.0MPa, H2/oil ratio: 550ml/ml)
80
4.3.2 Impact of H2S removal and LHSV ratio on cetane index
Results for cetane index (CI) improvements in the single and the two-stage
hydrogenation processes are shown in Table 4.3. Cetane index was calculated from the
ASTM D976 correlation in equation 4.3. Significant differences in the cetane indices
from the single and two-stage processes were observed at higher reaction temperatures
(i.e.380 -390 oC), where about 7-10 % increments were observed for the two-stage
process. At the lower temperatures (340-350 oC), only 2-3 % increase in the CI occurred
in the two-stage unit. The highest cetane index in the single-stage process was 44 at 365
oC whereas for the two-stage process, a higher temperature of 380 oC was required to
obtain the best cetane value of 46.
It is also observed from Table 4.3 that further increase in reactor temperature
above 380 oC (in the two-stage) and 365 oC (in the single-stage) had limitations for
additional increase in CI and this is because these temperatures are close to the
equilibrium temperature for hydrogenation of aromatics and consequently, the cetane
improvement. Figure 4.9 shows the impact of the H2S removal on the overall cetane
index improvement in the two-stage process at the different reaction temperatures and
LHSV ratios. Just like hydrogenation of aromatics the best CI occurred at 380 oC at the
LHSV ratio of 1.5:1.0.
4.3.3 Impact of H2S on HDS and HDN
The impact of removal of hydrogen sulfide on HDS in the two-stage compared to
the single-stage is shown in Figure 4.10. Similar to the HDA activity, the relative gain in
rate constants decreased with increasing reaction temperature. This is because majority
of hydrogen sulfide is produced in the stage I of the two-stage unit and adsorption of
81
Table 4.3: Cetane index improvement in single and two-stage processes at a pressure of 11.0 MPa
Cetane index, CI
Temperature [oC] Single -stage Two-stage
350 42 ± 0.8 43 ± 0.8
365 44 ± 0.9 44 ± 0.9
380 43 ± 0.9 46± 0.9
390 41 ± 0.8 45 ± 0.9
82
384042444648
350
365
380
390
Tem
pera
ture
[o C
]
Cetane Index,CI
Fig
ure
4.9:
Eff
ect
of H
2S r
emov
al a
nd L
HSV
rat
io o
n th
e ov
eral
l cet
ane
inde
x in
the
tw
o-st
age
proc
ess
(P
ress
ure:
11.
0MP
a, H
2/oi
l rat
io:
550m
l/ml)
LH
SV (
h-1)
1.5:
1
1:1
1:1.
5
Fee
d C
I =
36
83
hydrogen sulfide on the active sites of the catalysts decreases with increasing
temperature. Consequently, the rate of sulfur removal increases with further increase in
temperature in both stages leading to a net decrease in the overall gain. This observation
is consistent with other studies in literature (Kabe et.al., 1999; Botchwey et.al., 2003).
The positive gain in the rate constants implies higher HDS activity in the two-stage than
the single- stage. HDS activity at LHSV ratios of 1:1 and 1.5:1 were similar, however, at
lower temperatures (350-365 oC), higher HDS activity were observed at the reaction
conditions where the LHSV ratio was 1:1. From this results, it can be inferred that equal
catalyst loadings or a higher amount of NiMo catalyst loading in stage I will maximize
HDS activity in the two-stage process so long as there is an inter-stage removal of H2S.
In contrast to the HDA activity, inter-stage removal of H2S was significant to HDS at all
temperatures.
Unlike HDA and HDS very high HDN activities (95-100 % conversion) were
observed for both the single and two-stage processes within the temperature range
studied (350-390 oC). This implies negligible hydrogen sulfide inhibition for HDN.
Satterfield et.al., 1981 and Landau et.al., 1996 have also reported negligible hydrogen
sulfide inhibition on HDN reactions.
84
02468
350
365
380
390
Tem
pera
ture
[o C]
Relative gain in rate constant, ks [%]1:
11.
5:1
1:1.
5
Fig
ure
4.10
: Im
pact
of
H2S
inhi
biti
on o
n H
DS
in t
he t
wo-
stag
e pr
oces
s. (
Pre
ssur
e: 1
1.0M
Pa;
H2/
oil r
atio
: 55
0ml/m
l)
85
4.4. Kinetic studies
The purpose of this section of the research was to derive mathematical models
describing the rate of hydrotreating and mild hydrocracking (MHC) during light gas oil
upgrading, as a function of the active components of the system such as the
concentration and temperature. The Langmuir-Hinshelwood (L-H) rate of reaction
equation was used to kinetically model the hydrogenation and hydrodesulfurization data.
The L-H model was selected over that of the power law so as to account for any
hydrogen sulfide inhibition. However, the pseudo-first order power law was used to
model the MHC kinetics.
The experiments for the kinetic studies were performed by varying temperature
from 350-390 oC at the optimum pressure of 11.0 MPa. The LGO blend from Athabasca
bitumen was used as feedstock for the experiments. Simulated distillation data from of
ALGO, HLGO and VLGO hydrotreating were used to develop the MHC kinetics. It may
be noted that the catalyst packing and the experimental conditions were chosen such a
way to eliminate mass transfer resistances (see Section 3.4.1 and Bej et.al., 2001)
4.4.1 Single-stage kinetics with NiMo/Al2O3
Kinetic analysis of the single-stage hydrotreating process is divided into three
main sections: Kinetics of HAD; Kinetics of HDS and MHC kinetics.
4.4.1.1 Kinetics of HDA
It has been well established that hydrogenation of aromatics is an equilibrium
reaction which is shifted in favor of aromatics with increasing temperature:
Aromatics nH2 + Saturates (4.7)
86
Using the Langmuir-Hinshelwood rate of reaction equation to model the HDA kinetics:
++−
=−=−SHSHAA
AHRAHHAFAA PKCK
CkCPKKk
dt
dCr
22
22
1 (4.8)
where -rA, kF and kR are the rate of reaction, forward and reverse rate constants,
respectively, KH, KA and KH2S are the equilibrium adsorption constants of hydrogen,
aromatics and hydrogen sulfide, respectively. CA and CAH are the product concentration
of aromatics and the saturated species, respectively. PH2S and PH2 are partial pressures of
hydrogen sulfide and hydrogen gas, respectively and t is the residence time.
Analysis of the hydrotreating data showed negligible equilibrium effects leading
to the following assumptions for developing the final kinetic model:
• The surface reaction was rate limiting
• Reaction is pseudo first order in the forward reaction.
• The reaction occurred in a plug flow regime with negligible diffusion and mass
transfer effects (HDA is reaction- controlled)
• Reaction occurred in excess amounts of hydrogen at constant partial pressure
• Hydrogenation is inhibited by hydrogen sulfide which is produced from the HDS
process and H2S is an ideal gas. The partial pressure of hydrogen sulfide is
calculated from the ideal gas law equation:
)(22
2 spsoSHSH
SH CCbRTCRTV
nP −=== (4.9)
where PH2S is the partial pressure of H2S, nH2S is the number of moles, R is the universal
gas constant, T is temperature, Cso and Csp are the sulfur concentrations in the feed and
products, b is a constant and V is the volume of the solution.
87
The final equation is given by:
++=−=−
SIHSIHAIAI
AIHHIAIAIAIAI PKCK
CPKKk
dt
dCr
22
2
1 (4.10)
where kAI is the apparent rate constant and all other terms are as defined. The subscript I,
refers to the single-stage parameters.
The integral form of equation 4.10 was solved using MAPLE 6.0 software. Full
details of the calculation are given in Appendix D. The apparent kinetic parameters were
determined using non-linear least squares approach. Apparent activation energy and
heats of adsorption were also determined from the slopes of the curve fitting by plotting
the inverse of temperature against the logarithm of apparent kinetic and adsorption
equilibrium constants [ln (k, K) vs. 1/T] in Figure 4.11. The high correlation
coefficients obtained from the regression analysis indicated a good fit of the model to
the experimental data.
4.4.1.2 Kinetics of HDS
Unlike the reaction mechanism of hydrogenation of aromatics, HDS is known to
follow an irreversible pathway where:
Heteroatom sulfur species + Hydrogen Hydrocarbon + H2S (4.11)
The final rate expression for HDS is:
where rSI and kSI are the rate equation and apparent rate constant of HDS, respectively,
KS, KH2, KH2S are the adsorption equilibrium constants of sulfur, hydrogen and hydrogen
sulfide, respectively.
+++==−
SIHSIHHHISISI
SIHHISIsISIs PKPKCK
CPKKk
dt
dCr
I
222
2
1(4.12)
88
-4
-2
0
2
4
6
1.5 1.51 1.52 1.53 1.54 1.55 1.56 1.57 1.58
1000/T [K -1]
ln(k
,K)
lnk lnKA lnKH2S lnKH
Figure 4.11: Ar rhenius and Van’ t Hoff plot for single-stage HDA
89
The integral form of the rate expression in equation 4.12 is:
( ) ( )[ ]
Is
SIHSIHHHISI K
LambertWPKPKC
χ×++= 2221
(4.13)
where
( )SIHSIHHHI
SIHSIHHHI
HSIHISI
SIHSIHSOHHISOSOSISOHSIHIsI
sI
PKPK
PKPK
PKKk
PKCPKCCKCtPKKk
K
222
222
2
2222
1
1
)ln()ln()ln(
exp
)(++
�����
�
�
�����
�
�
++
���
����
�+++
−−
=χ (4.14)
And
2 3 4 5 6 73 8 125 54( ) ( )
2 3 4 5LambertW x x x x x x x o= − + − + − + (4.15)
The apparent kinetic parameters for HDS were also determined using nonlinear
least squares method. The activation energy and heats of adsorption were calculated
directly from the slopes of the Arrhenius and Van’ t Hoff plots in Figure 4.12. High
correlation coefficients greater than 0.985, were obtained from the regression analysis.
This indicated a good fit of the model to the experimental data.
4.4.1.3 MHC kinetics in ALGO, HLGO and VLGO
Hydrotreating and mild hydrocracking (MHC) are important catalytic processes
for producing high quality diesel fuels from petroleum feedstock. Compared to
hydrotreating however, the MHC mode of operation requires higher reactor temperatures
(Yui et.al., 1989). It also improves hydrogen consumption economy and minimizes
formation of undesirable lighter products (Satterfield, 1981). MHC has an advantage
over conventional hydrocracking in improving the cetane rating of diesel by ring
90
-6
-4
-2
0
2
4
1.5 1.52 1.54 1.56 1.58 1.6 1.621000/T [K -1]
ln(k
s,K
)
lnksI lnKHI lnKsI lnKH2S
Figure 4.12: Ar rhenius and Van’ t Hoff plots for single-stage HDS
91
opening of saturated rings thus converting naphthalene and branched-alkanes to lighter,
gasoline-range products.
The projection of the long term middle distillate shortage in the mid 1980s
spurred researchers and petroleum refiners to investigate mild hydrocracking of heavy
gas oil with the view to shifting the product slate toward increased middle distillate
production (Yui et.al., 1989). Middles distillates serve as a feed source for diesel fuel
production. However, most of the literatures on MHC by workers including Desai et.al
1985; Wilson et.al., 1987; Hill et.al., 2002, have focused mainly on the product yields,
properties and optimum conditions to produce a maximum distillate yield. Reports on
the kinetics of MHC are limited.
The main objective of this part of the research was to investigate the MHC
kinetics describing the results of the product yield in the three light gas oil feedstock
(ALGO, VLGO and HLGO). MHC was measured by the extent of heavy gas oil (345+
oC) conversion:
)(345
)(345)(345)345(
feedC
productsCfeedCCConversion
o
ooo
++−+=+ (4.16)
The feed and product samples were divided into three main fractions: Gasoline (G) (40–
205ºC), Diesel (D) (205-345 ºC) and the heavy gas oil fractions (H) (345 + ºC).
Conversion of the heavy gas oil fractions (H) was assumed to follow the parallel
reaction mechanism in scheme 1:
Scheme 1: Reaction pathway for conversion between 340-390 oC (H: 345+ oC; D: 205-345+oC; G: 40-205 oC)
H
D
G
k1
k2
92
Cracking of diesel to gasoline was assumed to be negligible and cracking of the heavy
gas oil fractions into gasoline and diesel was also assumed to obey pseudo first-order
kinetic and given by:
( )1 2H
H
dCk k C
dt= − + (4.17)
( )1 2H HO
k k tC C e
− += (4.18)
where CH and CHO are the product and feed mass fractions of the heavy has oil fraction,
respectively, k1 and k2 are the rate constants for cracking into diesel and gasoline,
respectively, and t is the reaction time.
Analysis of the MHC kinetic data in Table 4.4 shows that the activation energies
of gasoline production (k2) are much higher than those for diesel production (k1) which
is an indication that cracking to gasoline increases strongly as the temperature increases.
Similar results have also been reported by Yui et.al., 1989 when they investigated MHC
kinetics on bitumen-derived heavy gas oils. In terms of the activation energies of the
combined rate constants (k*=k1+k2), the ease of mild hydrocracking in the LGO
feedstock followed the order: HLGO > VLGO > ALGO. Thus, more cracking products
are obtained from the HLGO, followed by the VLGO and finally the ALGO.
93
Table 4.4: Kinetic parameters of MHC in ALGO, HLGO and VLGO
E [kJ/mol] ko R2
ALGO
k1 52 2.9 ×103 0.9906
k2 102 2.3 ×107 0.9972
k* = k1+k2 88 4.8 ×106 0.9869
HLGO
k1 41 2.6 ×102 0.9913
k2 62 9.4 ×103 0.9923
k* = k1+k2 49 2.0 ×103 0.9915
VLGO
k1 55 3.7 ×103 0.9937
k2 85 7.5 ×105 0.9942
k* = k1+k2 67 5.8 ×104 0.9946
94
4.4.2 Two-stage kinetic studies
This section of the study, discusses the development of the overall kinetics of
HDA and HDS in the two-stage process. Results from the kinetics studies are compared
to those from the single-stage process to determine the effect of H2S inhibition on the
hydrotreating activities.
4.4.2.1 Overall HDA and HDS kinetics
The two-stage process was performed to study the effect of H2S removal on
HDA and HDS in light gas oil from Athabasca bitumen. Prior to the stage II reactions,
H2S present in the products from stage I were completely removed by bubbling nitrogen
gas through the samples for at least two hours. The H2S produced in the stage II
reactions is due to the unreacted sulfur species from the stage I reaction effluents. The
rate equations describing the overall HDA and HDS kinetics were similar to those
developed in the single-stage process (see Appendix D for the HDA kinetic equations):
For HDA, the overall rate of reaction equation is:
+++=−=−
SIIHSIIHHHIIAIAII
AIIHHIIAIIAIIAIIAII PKPKICK
CPKKk
dt
dCr
222
2
1 (4.19)
Overall HDS rate of reaction equation is also defined by:
(4.20)
where PH2SII = b (CSI-CSII) and all parameters are as defined earlier. The
subscripts I and II refer to stage I and stage II, respectively.
The overall kinetic parameters were also determined using the non-linear least
square regression approach. The activation energies and heats of adsorption were
+++=−=−
SIIHSIIHHHIISIISII
SIIHHIISIIsIISIIsII PKPKCK
CPKKk
dt
dCr
2221
95
obtained directly from the slopes of the Arrhenius and Van’ t Hoff plots. High
correlations coefficients were also derived for the regression analyses which indicated
accurate predictions of the experimental data by the models. A summary of the single-
stage and the two-stage (overall) kinetic parameters for both HDA and HDS are shown
in Table 4.5.
4.4.2.2 Effect of H2S removal on HDA kinetics
The following observations were made from the kinetic parameters when
hydrogen sulfide was removed from the two-stage process:
The rate constants, kA derived from the single-stage hydrogenation activity were
lower (1.8-6.2×10-5s-1) than those from the two-stage hydrogenation data. Thus a faster
rate of reaction favored the two-stage process.
For the single-stage process, the equilibrium adsorption constants of aromatics
were lower than those for hydrogen sulfide whereas for the two-stage process, the
opposite was observed. Hence, the aromatic compounds in the single-stage process were
weakly adsorbed (low conversion) due to H2S inhibition while for the two-stage process,
the aromatic species were strongly adsorbed, leading to a higher HDA activity and
consequently, higher conversions.
The activation energy of hydrogenation in the single-stage process was 85 kJ/mol
but upon removal of hydrogen sulfide for the two-stage process, the activation energy
dropped to 67 kJ/mol. This implies that when HDA in the bitumen-derived light gas oil
is retarded by H2S, more energy will have to be provided to the system in order to
overcome the energy barrier and reduce the aromatics contents to lower levels below 10
vol % in compliance with the current diesel fuel specifications (US EPA, 1999).
96
Table 4.5: Summary of the apparent kinetic parameters of the overall kinetics studies in the single and two-stage processes (Temperature-340 -390 oC; Pressure-11.0 MPa; Total residence time-1.67 h) Kinetic
Parameters
lnkAI
lnkAII
lnKAI
lnKAII
lnKH2SI
lnKH2SII
lnKHI
lnKHII
Aromatics Hydrogenation(HDA) EA and ∆H [kJ/mol]
85 ± 4.2
67 ± 3.3
33 ± 1.7
59 ± 3.0
39 ± 2.0
24 ± 1.2
7.6 ± 0.38
7 ± 0.36
ln (kAo,Ko)
14 ± 0.70
10.6 ± 0.53
-4.8 ± 0.24
-8.2 ± 0.41
-3.5 ± 0.20
-5.4 ± 0.27
-2.3 ± 0.14
-4.8 ± 0.24
R2
0.9967
0.9854
0.9995
0.9963
0.9729
0.9999
0.9833
0.9988
Hydrodesulfur ization (HDS)
lnksI
lnksI I
lnKSI
lnKSII
lnKH2SI
lnKH2SII
lnKHI
lnKHII
ES and ∆H [kJ/mol]
55 ± 2.8
22 ± 1.1
44 ± 2.2
67 ± 3.4
50 ± 2.5
24 ± 1.2
114 ± 5.7
-33 ± 1.6
ln (kso,Ko)
5.2 ± 0.26
0.78 ± 0.04
-5.5 ± 0.28
-9.0 ± 0.45
-8.0 ± 0.40
-2.4 ± 0.12
-19 ± 0.98
4.4 ± 0.22
R2
0.9963
0.9972
0.9868
0.9973
0.9941
0.9943
0.9863
0.9984
97
4.4.2.3 Effect of H2S removal HDS kinetics
For the analysis of the single-stage overall kinetics, the activation energy of HDS
was 55 kJ/mol. The heats of adsorption for sulfur and hydrogen sulfide were 44 and 50
kJ/mol, respectively. Therefore, H2S was strongly adsorbed on to the active sites of the
catalyst as compared to adsorption of the sulfur species. Furthermore, the adsorption
equilibrium constants for H2S were higher than those for total sulfur.The results suggest
that in the HDS of light gas oil, the reaction was retarded by the H2S produced in the
reaction. This result is consistent with the studies by Kabe et.al., 1999 who reported a
high adsorption constant of H2S than those of dibenzothiophene compounds.
In the case of the two-stage kinetic analysis, the activation energy was 23 kJ/mol
while the heats of adsorption were 67 and 24 kJ/mol for the total sulfur and hydrogen
sulfide, respectively. Unlike the single-stage, the adsorption equilibrium constants of
total sulfur were significantly higher than H2S indicating that H2S was weakly adsorbed.
The inter-stage removal of H2S from the two-stage process greatly enhanced the HDS
activity as activation energy decreased from 55 kJ/mol in the single-stage to 23 kJ/mol
in the two-stage.
4.4.3 Exper imental versus model predictions The best kinetic parameters for predicting the experimental results for HDA,
HDS and MHC were selected based on the sum of square errors (SSE) approach.
2exp )( ip
i yySSE −= ∑ (4.21)
where SSE is the divergence, yip is the model prediction and yexp is the experimental
results.
98
The criterion for this method was to minimize the differences between the experimental
and the predicted results. Figures 4.13 and 4.14 compare the experimental and correlated
product concentrations of aromatics and sulfur, respectively for three sets of
hydrotreating data (same operating conditions). Figure 4.15 also compares the calculated
and experimental concentrations of the heavy gas oil fraction, (H) (345+oC) in the three
LGO feedstock. As can be seen, the kinetic models predicted with reasonable accuracy
the experimental results (with R2 ≥ 0.99) over the entire temperature range.
99
0369121518
03
69
1215
18
Exp
erim
enta
l CA [
%]
Correlated CA [%]
1.5:
11:
11:
1.5
F
igur
e 4.
13:
Cor
rela
ted
vers
us e
xper
imen
tal c
once
ntra
tion
s of
tot
al a
rom
atic
s, C
A a
t th
e di
ffer
ent
LH
SV r
atio
s
R2 =0
.989
9
100
0
50
100
150
200
250
300
0 50 100 150 200 250 300Exper imental Cs [wppm]
Cor
rela
ted
Cs [
wpp
m]
1.5:1.0 1:1 1:1.5
Figure 4.14: Cor related vs. exper imental concentrations of product sulfur concentrations (Cs) at the different LHSV ratios
R2=0.9979
101
0
10
20
30
40
50
0 10 20 30 40 50
Exper imental CH [%]
Cor
rela
ted
CH [
%]
VLGO ALGO HLGO
Figure 4.15: Cor related vs. exper imental concentrations of the heavy gas oil (345+oC) fractions from the mild hydrocracking data (simulated distillation) in VLGO, ALGO and HLGO
R2=0.9985
102
5.0 CONCLUSIONS
Interaction between temperature and pressure is the most significant factor
affecting HDA while HDS and HDN are highly influenced by interaction between
temperature and space velocity. The optimal conditions of HDA for maximum
conversion of aromatics were found at a temperature, pressure and LHSV combination
of 379 oC, 11.0 MPa and 0.6 h-1 respectively. At these conditions, the highest conversion
of 63 % could be attained. HDS and HDN conversions at these conditions were 98.5 and
99.7 %, respectively.
Hydrogenation of monoaromatics is the key step for reducing the total aromatics
content of light gas oils. The ease of hydrogenation of total aromatics in the LGO
feedstock was observed to follow the general order: VLGO > ALGO > HLGO. Studies
on MHC indicated a net increase in gasoline with a corresponding decrease in diesel
during cracking. More cracking products were produced from the HLGO feed.
HDA, cetane rating and HDS processes were inhibited by hydrogen sulfide
during hydrotreating in the single-stage reactor. However, with the two-stage process
where hydrogen sulfide was removed inter-stage, significant improvement in
hydrogenation, cetane and sulfur removal were observed.
HDA and HDS in LGO feedstock from Athabasca bitumen can be described by
the Langmuir-Hinshelwood kinetic models. Mild hydrocracking was best described by a
pseudo first-order parallel model.
103
6.0 RECOMMENDATIONS
From the experimental results obtained in this research, the following
recommendations can be made:
1. Further studies could be carried out to determine the inhibition effects of
ammonia (NH3) on aromatics hydrogenation.
2. Future experiments could be also be carried out to study the hydrogenation
activity of ring-opening catalysts such as the noble-metal catalysts for
improving the diesel quality of light gas oil fractions from Athabasca oil
sands.
3. For scaling up of the process, the design and development of two-fixed
reactors arranged in series with inter-stage hydrogen-sulfide removal should
be examined.
104
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APPENDIX
115
Appendix A: Exper imental calibration
A.1 Calibration of mass flow meter
The mass flow controller was calibrated for hydrogen flow at the experimental
operating conditions using a bubble flow meter connected to the exit of the backpressure
regulator. The flow rates measured at atmospheric conditions were standardized using
equation A.1.
a
o
a
o
o VT
T
P
PV
a
= (A.1)
Where V is the flow rate in ml/hr, T is temperature, P is pressure, the superscripts ‘o’
and a represent standard conditions normal operating conditions, respectively.
Figure A.1 shows the calibration curve of the mass flow controller.
A.2 Reactor temperature calibration
The reactor was temperature calibrated at the same conditions as used in the
actual experimental runs. Temperature was varied from 150 to 420 oC while maintaining
the reaction pressure and hydrogen-to-oil ratio constant at of 9.6 MPa and 550 ml/ml,
respectively. The corresponding reactor temperature was measured using a single
thermocouple inserted just below the catalyst bed. The thermocouple was then moved
every 1 cm up the catalyst bed to measure the temperature along the reactor bed. Profile
for the variation of the temperature along the reactor bed is shown in Figure A.2 and
Figure A.3 shows the calibration curve for the temperature controller.
116
y = 5.2675x - 0.1002
R2 = 0.9988
0
50
100
150
200
250
300
350
400
0 10 20 30 40 50 60 70
Set point [%]
Hyd
roge
n fl
ow r
ate
at S
TP
[m
l/min
]
Figure A.1: Calibration curve for mass flow meter
117
100
150
200
250
300
350
400
450
0.0 2.0 4.0 6.0 8.0 10.0 12.0 14.0
Reactor Length [cm]
Rea
ctor
Tem
pera
ture
[o C
]
150 oC
200 oC
250 oC
300 oC
320 oC
340 oC
355 oC
370 oC
385 oC
400 oC
420 oC
Catalyst bed
Figure A.2: Temperature distr ibution along the axial length of the reactor
11
8
y =
0.99
27x
- 2.
1437
R2 =
1
100
150
200
250
300
350
400 10
015
020
025
030
035
040
0
Set
poin
t T
empe
ratu
re [
o C]
Average Catalyst Bed Temperature [oC]
Fig
ure
A.3
: T
empe
ratu
re c
alib
rati
on c
urve
of
reac
tor
119
Appendix B: Feed and product analyses
B.1 Analysis for aromatics contents
The aromatics contents of the feed and products were determined using 13C-
NMR spectrometry and Supercritical fluid chromatography. The former was used to
determine the total aromatics content (aromaticity) while the latter was used to
determine the individual concentrations of the mono-, di- and polyaromatics.
B.1.1 13C-NMR spectrometry
Aromaticity is the mole per cent of carbon (%CA) in a sample that is present as
part of an aromatic ring structure. Feed and product aromaticity were measured directly
from the carbon-13 nuclear magnetic resonance (13C –NMR) spectroscopy. The spectra
were obtained in the Fourier Transform (FT) mode operating at a frequency of 500
MHz. The instrumental conditions were: pulse delay of 4 s, sweep width of 27.7 kHz
and inverse gated proton coupling. Overall time for each sample was one hour, 30
minutes for 2000 scans. Deuterated chloroform, CDCl3 was used to dilute the samples.
Figure B.1 shows a typical spectrum for 13C-NMR with two distinct zones
separated by the solvent bar. These are the total aromatics found between 120-150 ppm
and total saturated hydrocarbons between 0-50 ppm. The total aromaticity of each
sample is measured directly from the spectra by finding the percentage of total aromatics
from the equation:
100×+
=satra
ar
II
ICar % (B.1)
Where Iar is the integral of total aromatics; Isat = integral of total saturates; Car is the
aromatics content.
12
0
Fig
ure
B.1
: Sa
mpl
e 13
C-N
MR
spe
ctra
for
a h
ydro
trea
ted
sam
ple
121
B.1.2 Supercr itical fluid chromatography (SFC)
The individual concentrations of the mono, di and polyaromatics contents were
determined using the supercritical fluid chromatography technique. This is a relatively
recent chromatographic technique which is an adaptation of the high performance liquid
chromatography (HPLC). The major modification is the replacement of the liquid
mobile phase with a supercritical fluid mobile phase (carbon dioxide, CO2). A
supercritical fluid chromatograph consists of a liquid supply, usually CO2, a pump, the
column in a thermostat-controlled oven, a restrictor to maintain the high pressure in the
column and a detector.
The mobile phase is initially pumped as liquid and brought into the supercritical
region by heating it above its critical temperature before it enters the analytical column.
It then passes through an injection valve where the sample is introduced into the
supercritical stream and then into the analytical column. The mixture of the sample and
the mobile phase is maintained supercritical as it passes through the column and into the
detector by a pressure restrictor placed either after the detector or at the end of the
column. The column contains a highly viscous liquid (stationary phase) into which the
analytes can be temporarily adsorbed and then released based on their chemical
structure; the monoaromatics are eluted first followed by the di- and the polyaromatics.
Part of the theory of separation in SFC is based on the density of the supercritical
fluid which corresponds to solvating power. As pressure in the system is increased, the
supercritical fluid density increases with a corresponding increase in the solvating
power. Therefore as the density of the mobile phase is increased, components retained in
the column can be made to elute.
122
B.2 Total sulfur analysis
Total sulfur in the feed and sample products was analyzed by
combustion/fluorescence technique as provided by the ASTM D5453 method. The
hydrocarbon sample is injected into a sample boat and then inserted into a high-
temperature combustion tube where sulfur is oxidized to sulfur dioxide in an oxygen rich
atmosphere. Any water produced during sample combustion is removed and the sample
combustion gases are next exposed to Ultraviolet (UV) light. The SO2 absorbs the
energy from the UV light and is converted to excited SO2* . The fluorescence emitted
from the excited SO2* as it returns to its stable state SO2 is detected by a photomultiplier
tube and the resulting signal is a measure of the sulfur contained in the sample. The
sulfur content of the test specimen in parts per million (ppm) is calculated as:
( )
KgMS
YIppmCS ××
−=*
)( (B.2)
Where Cs is the concentration of sulfur, I is the average integrated detector response for
test specimen solution, counts; Y is the y-intercept of standard curve, counts; S is the
slope of standard curve, counts/mg, M* is the mass of test specimen solution injected
and Kg of the gravimetric dilution factor, mass of test specimen/mass of test specimen
and solvent, g/g.
B.3 Total nitrogen analysis
The feed and sample products are injected into a sample boat. A helium or argon
carrier gas then sweeps the sample into a pyrolysis tube. The nitrogen in the sample is
then oxidized to nitric oxide (NO) in an oxygen chamber. The oxides of nitrogen are
contacted with ozone (O3) which converts NO to NO2. As the metastable species decays
123
a photon of light is emitted and detected by the photomultiplier. The
chemiluminescence’s emission is specific for nitrogen and is proportional to the amount
of nitrogen on the original sample. The nitrogen content, CN in parts per million is
evaluated as:
( )KgMS
YIppmCN ××
−=*
)( (B.3)
Where CN is the concentration of nitrogen, I is the average integrated detector response
for test specimen solution, counts; Y is the y-intercept of standard curve, counts; S is the
slope of standard curve, counts/mg, M* is the mass of test specimen solution injected
and Kg of the gravimetric dilution factor, mass of test specimen/mass of test specimen
and solvent, g/g.
B.4 Simulated distillation
The simulated distillation chromatography analysis method is a substitute for
conventional distillation methods to estimate parameters (boiling temperature
distribution) for large scale petroleum refining process. A Varian Model CP 3800 Gas
Chromatography (especially configured for simulated distillation) coupled to a Varian
CP 8400 auto sampler was used for the boiling point distribution analysis. The simulated
distillation chromatography is used to determine the boiling range distribution of crude
petroleum and various petroleum fractions and products by assigning the boiling
temperatures as a function of retention time. The temperatures at which specific
percentages of total sample elutes from the column is then measured.
124
B.5 Catalyst (NiW/Al2O3) character ization
BET surface area, pore volume and size of the fresh and spent catalysts were
determined using an automated gas (N2) adsorption analyzer ASAP 2000 (Micrometrics)
with pure nitrogen gas (99.9 % pure). About 0.05 g of sample was used and before each
analysis, the catalyst sample was evacuated at 200 oC for 4h in a vacuum to remove all
adsorbed moisture from the catalyst surface. Prior to surface area/porosity
determinations, all spent catalysts were thoroughly washed with hexane solution to
remove volatiles as well as gas oil present on the surface and in the pores of the
catalysts. Cleaned catalysts were then dried in an oven at 120 °C for 12 hours.
BET surface area, pore volume and pore diameter of lab-prepared NiW/Al2O3
catalyst were 174 m2/g, 0.495 cc/g and 114 Å, respectively. In the case of the
commercial NiMo/Al2O3 catalyst, the BET surface area, pore volume and pore diameter
were reported to be 169 m2/g, 0.412 cc/g and 97.8 Å, respectively.
High resolution transmission electron microscopy (TEM) analyses of sulfided
NiW/Al2O3 catalysts was carried out with a Philips CM20 electron microscope with a
LaB6 filament as a source of electron and operated at 200kV. The purpose of the TEM
analysis was to determine the dispersion of the WS2, which is the active phase of the
hydrotreating catalysts in its working state. For the TEM analysis, the sample was
cleaned sample was powdered and a small amount of the powdered sample was then
sprinkled on a piece of parafilm with a droplet of water. A 400 mesh carbon coated grid
was floated on the droplet of water and then picked. The retained droplet of material was
allowed to air dry on the grid and then mounted on a specimen holder where the analysis
was carried out.
125
Figure B.2 and Figure B.3 show the TEM micrographs of the NiMo/Al2O3 and
NiW/Al2O3 catalysts, respectively. The lattice images of the MoS2 and WS2 slabs are
indicated by the solid black lines in both figures. The sulfided NiMo/Al2O3 catalyst
showed higher dispersion of MoS2 slab compared to that of WS2 in case of sulfided
NiW/Al2O3 catalyst.
Figure B.2: TEM micrograph of the sulfided NiMo/Al2O3 catalyst
126
Figure B.3: TEM micrograph of the sulfided NiW/Al2O3 catalyst
WS2
127
Appendix C: Log sheets
C.1 Data recording
Table C.1 is an example of the log sheets used to record data and monitor
experiments. The experiments were monitored for pump performance, temperature,
pressure, space velocity, flow rates and hydrogen-to-oil ratio.
12
8
Tab
le C
.1:
Sam
ple
of t
he d
ata
reco
rdin
g sh
eet
Rea
ctor
#:
5
Dat
e:
2/05
/200
4
Fee
d T
ype:
L
GO
B le
nd
C
atal
yst
Typ
e: N
iW/A
l 2O
3
Dat
e d.
m:y
r
Tim
e
(h)
TO
S
(h
)
H2
syst
em p
ress
ure
(psi
g)
T
ank.
PG
I P
G3
PG
4
Tem
pera
ture
(o C
) Fu
rnac
e
Rxt
Pum
p se
t po
int
(%
)
Oil
W
eigh
t
(g)
Flow
ra
te
(g
/h)
LH
SV
(h-1
)
H2
MFM
(ml/
ml)
Rem
arks
129
Appendix D: Exper imental calculations and mass balance closure.
D.1 Equations used for calculating the HDA kinetic parameters
The Langmuir-Hinshelwood rate equation model was used to describe the
kinetics of hydrogenation of aromatics. This is given by:
SiHSiHAiAi
AiHHiAiAiAiAi PKCK
CPKKk
dt
dCr
22
2
1 ++=−=− (D.1)
Where -rA, kA are the rate of reaction, and the forward rate constants, respectively, KH,
KA and KH2S are the equilibrium adsorption constants of hydrogen, aromatics and
hydrogen sulfide, respectively. CA is the product concentration of aromatics and the
saturated species, respectively. PH2S and PH are partial pressures of hydrogen sulfide and
hydrogen gas, respectively and t is the residence time. The subscript, i, refers to either
the single-stage or two-stage kinetic parameters.
Partial pressure of hydrogen sulfide is assumed to be an ideal gas and calculated by:
)(22
2 spsoSHSH
SH CCbRTCRTV
nP −=== (D.2)
2 2
2 22 2
2 2
ln ln( )
exp( ))1
(1 )1
A
AO A Ao AO H S H SA H H
A H H
H S H SH S H S
H S H S
AA
C K C C K PkK K P t
kK K PK
K PK P LambertW
K P
CK
+ + − + + + = (D.3)
Where
2 3 4 5 6 73 8 125 54( ) ( )
2 3 4 5LambertW x x x x x x x o= − + − + − + (D.4)
130
And
2 2
2 2
2 2
ln ln( )
exp( ))1
1
A
AO A Ao AO H S H SA H H
A H H
H S H S
H S H S
C K C C K PkK K P t
kK K PK
K PX
K P
+ + − + = +
(D.5)
D.2 Mass Balance Calculations
The mass balance closure for aromatics, sulfur, and nitrogen and the hydrocrack
materials were calculated using the following steady-state equations:
For HDA:
Cso + Cmo + Cdo +Cpo = Cs + Cm +Cd +Cp (D.6)
Where Cso, Cmo, Cdo and Cpo represent the percentage of saturates, mono-, di-, and
polyaromatic compounds in the feed, respectively Cs, Cm, Cd and Cp are the percentage
saturates, mono, di- and polyaromatics in the hydrotreated products, respectively.
Overall material balance closure was 99.5 %.
For HDS:
CSF = CSP + CH2S (D.7)
Where CSF and CSP represent the concentrations of sulfur in the feed, and products,
respectively and CH2S is the concentration of hydrogen sulfide gas produced during
reaction. The overall material balance closure was 97 %.
131
For HDN:
CNF = CNP + CNH3 (D.8)
Where CNF and CNP represent the concentrations of nitrogen in the feed and products,
respectively and CNH3 is the concentration of ammonia produced during HDN reactions.
For MHC:
CHO + CDO + CGO = CH + CD + CG (D.9)
Where CHO, CDO and CGO are the percentage amounts of heavy gas oil, diesel and
gasoline in the feed, respectively. CH, CD and CG are the percentages of heavy gas oil,
diesel and gasoline in the products, respectively. The overall mass balance closure,
ignoring gases mixed in the liquid products, ranged from 97-98.5 % for all the light gas
oil feedstock.
132
Appendix E: Exper imental results Table E.1: Total aromatics, sulfur and nitrogen concentrations after the single-stage hydrotreating with commercial NiMo/Al2O3 catalyst
Temp
[oC]
Pressure
[MPa]
LHSV
[h-1]
13C-NMR
[%]
Sulfur
[wppm]
Nitrogen
[wppm]
Feed - - 17.1 17420 461
365 6.9 1.25 5.2 305 20
390 9.6 1.25 4.0 201 10
365 9.6 1.25 5.4 1163 63
350 9.6 1.25 4.9 403 11
380 12.4 1.70 5.9 430 16
350 8.3 1.25 7.7 1310 96
380 8.3 1.70 7.5 286 21
365 11.0 1.70 6.9 1385 63
350 11.0 2.00 4.6 732 19
380 9.6 0.80 4.9 393 20
350 8.3 1.00 6.9 346 14
380 8.3 0.80 6.5 191 9
350 11.0 0.80 6.6 617 13
380 11.0 0.50 3.6 298 10
365 9.6 1.25 4.1 261 10
365 9.6 1.25 4.5 258 <10
365 9.6 1.25 4.4 262 <10
365 9.6 1.25 3.9 260 <10
365 9.6 1.25 4.0 263 <10
13
3
Tab
le E
.2:
Aro
mat
ics,
sul
fur
and
nitr
ogen
con
cent
rati
ons
in t
he L
GO
ble
nd a
fter
sin
gle-
stag
e hy
drot
reat
ing
SFC
Aro
mat
ic c
onte
nts
[%]
Sam
ple
ID
Tem
p
[o C]
Pre
ssur
e
[MP
a]
LH
SV
[h-1
]
Satu
rate
s
[%]
Mon
o-
Di-
P
oly-
T
otal
Sulf
ur
[wpp
m]
Nit
roge
n
[wpp
m]
Hyd
rotr
eati
ng o
ver
NiM
o/A
l 2O
3
34
0 11
.0
0.6
75.9
0 20
.30
3.60
0.
150
24.0
8 87
1 16
35
0 11
.0
0.6
81.7
0 16
.10
2.20
0.
030
18.3
2 57
6 <1
0
36
5 11
.0
0.6
87.9
0 10
.70
1.30
0.
004
12.0
5 29
7 <1
0
38
0 11
.0
0.6
91.4
0 7.
70
0.87
0.
004
8.55
22
2 <1
0
39
0 11
.0
0.6
95.6
0 4.
60
0.40
0.
001
4.98
27
1 <1
0
Hyd
rotr
eati
ng o
ver
NiW
/Al 2
O3
34
0 11
.0
0.6
68.0
5 6.
11
3.60
0.
830
34.7
1 21
05
55
35
0 11
.0
0.6
69.8
2 5.
26
0.83
0.
530
33.0
3 10
49
25
36
5 11
.0
0.6
76.7
3 3.
09
0.53
0.
090
26.8
2 94
9 <1
0
38
0 11
.0
0.6
83.4
0 1.
93
0.09
0.
040
20.1
5 75
8 <1
0
39
0 11
.0
0.6
85.7
4 1.
42
0.04
0.
010
17.8
4 74
8 <1
0
134
Table E.3: Simulated distillation data obtained for the feed character ization study of the different light gas oil feedstock.
Fraction ALGO LGOB HLGO VLGO
Mass, wt % Boiling point [oC]
IBP 125 130 133 254
5 180 191 179 288
10 212 219 197 301
15 228 236 211 311
20 240 250 224 318
25 252 262 236 325
30 261 273 248 331
35 269 283 258 337
40 278 292 269 342
45 287 302 280 347
50 295 310 291 353
55 303 319 301 358
60 311 327 311 364
65 319 336 321 370
70 329 345 332 376
75 339 355 343 383
80 350 366 356 391
85 364 378 369 401
90 381 395 386 415
95 409 420 410 436
100 476 484 472 492
13
5
Tab
le E
.4:
Mild
hyd
rocr
acki
ng d
ata
of L
GO
typ
es a
t a
pres
sure
of
11.0
MP
a
and
LH
SV o
f 0.
6 h-1
Fee
d ty
pe
Tem
p.
[o C]
Gas
olin
e [w
t %
] D
iese
l [w
t %
] H
GO
[w
t %
] 34
0 19
63
18
350
20
63
17
365
21
64
15
380
22
65
13
AL
GO
390
23
68
9
340
27
55
18
350
28
55
17
365
26
57
17
380
27
57
16
HL
GO
390
27
60
14
340
0.9
57
48
350
0.7
51
42
365
3 59
44
380
4 59
36
VL
GO
390
7 59
35
13
6
Tab
le E
.5:
Ove
rall
arom
atic
s, s
ulfu
r an
d ni
trog
en c
once
ntra
tion
s in
the
LG
O b
lend
aft
er t
he t
wo-
stag
e hy
drot
reat
ing
proc
ess
SFC
Aro
mat
ic c
onte
nts
[%]
Sam
ple
ID
Tem
p
[o C]
Pre
ssur
e
[MP
a]
LH
SV
[h-1
]
Satu
rate
s
[%]
Mon
o-
Di-
P
oly-
T
otal
Sulf
ur
[wpp
m]
Nit
roge
n
[wpp
m]
Feed
-
- -
63.3
20
.7
12.2
3.
6 36
.5
1742
0 46
1
SS-3
50-1
35
0 11
.0
1.5
83.2
14
.6
2.2
0.03
16
.8
262
<10
SS-3
65-1
36
5 11
.0
1.5
87.4
10
.9
1.7
0.03
12
.6
690
<10
SS-3
80-1
38
0 11
.0
1.5
92.7
6.
3 1.
0 0.
02
7.3
233
<10
SS-3
90-1
39
0 11
.0
1.5
92.2
6.
7 1.
1 0.
05
7.8
167
<10
SS-3
50-2
35
0 11
.0
1.2
86.7
11
.9
1.4
0.02
13
.3
78
<10
SS-3
65-2
36
5 11
.0
1.2
91.2
8.
0 0.
9 0.
00
8.9
97
<10
SS-3
80-2
38
0 11
.0
1.2
92.9
6.
4 0.
7 0.
00
7.1
88
<10
SS-3
90-2
39
0 11
.0
1.2
92.3
7.
0 0.
7 0.
00
7.7
79
<10
SS-3
50-3
35
0 11
.0
1.0
87.2
11
.5
1.3
0.00
12
.8
109
<10
SS-3
65-3
36
5 11
.0
1.0
91.4
7.
7 0.
9 0.
00
8.6
95
<10
SS-3
80-3
38
0 11
.0
1.0
93.0
6.
2 0.
8 0.
00
7.0
80
<10
SS-3
90-3
39
0 11
.0
1.0
93.9
5.
8 0.
3 0.
00
6.4
68
<10
